Efficient treatment of wastewater using electrochemical cell

ABSTRACT

An efficient method and system for the electrochemical treatment of waste water comprising organic and/or inorganic pollutants is disclosed. The system comprises at least first and second solid polymer electrolyte electrolytic cell stacks in which each cell comprises a solid polymer, proton exchange membrane electrolyte operating without catholyte or other supporting electrolyte. The first and second stacks differ either in construction or operating condition. The cell stack design and operating conditions chosen provide for significantly greater operating efficiency.

FIELD OF THE INVENTION

The present invention relates to methods and systems for the electrochemical treatment of waste water. In particular, it relates to methods and systems for the removal of organic pollutants and oxidation of inorganic compounds using solid polymer membrane electrolyte electrochemical cells.

BACKGROUND

There is substantial growth in the demand for new wastewater treatment technologies that is being driven by population growth and increasing volumes of wastewater produced, tighter wastewater quality regulations, increasing cost of clean water and water shortages, awareness for the protection of clean water sources and replacement of aging wastewater systems. Industries are specifically being forced both by tougher discharge standards and cost pressures to eliminate their recalcitrant wastewater pollutants prior to discharge, and to adopt on-site water reuse and recycling systems to avoid rising water supply and effluent discharge costs. The requirement is for cost-effective, sustainable water treatment technology that does not require the addition of chemicals and does not produce secondary pollution, is compliant with stringent water quality standards, and has minimal operational and maintenance requirements.

Industrial wastewater can contain organic compounds, many of which are toxic, persistent and resist conventional biological and chemical wastewater treatment. The best approach to treat recalcitrant wastewater is by non-chemical oxidation techniques that can mineralize the pollutants and reduce the organic load and toxicity of the waste, such as electrochemical oxidation. Electrochemical oxidation is sustainable, safe and has a high treatment efficacy eliminating a wide variety of pollutants such as persistent organic pollutants, dioxins, nitrogen species (e g ammonia), pharmaceuticals, pathogens, microorganisms, a majority of priority pollutants and pesticides. There are two main approaches to electro-oxidation of pollutants in wastewater. The first is to oxidize pollutants by indirect electrolysis, generating a redox reagent in situ as a chemical reactant. The mediator can be a metallic redox couple or a chemical reagent (e.g. chlorine, ozone, peroxides). These processes require the addition of a large amount of chemicals and/or feed oxygen, and produce secondary pollution leading to additional costs for the disposal of the treated wastewater and operation and maintenance of the process. The second approach is to use direct electrochemical oxidation, where the organic pollutants are oxidized on the anode surface.

A variety of cell configurations that include flow-through parallel plates, divided chambers, packed bed electrodes, stacked discs, concentric cylinders, moving bed electrodes and filter-press have been developed for direct electrochemical wastewater treatment. However, common to all these electrochemical cell configurations is poor operational efficiency leading to high energy consumption. The wastewater is utilized as electrolyte, and in the case of divided cells, both anolyte and catholyte. Due to very low ionic conductivity of wastewater though, the addition of a supporting electrolyte is required to improve the cell efficiency and obtain reasonable cell voltages. This generally results in salt, base and/or acid concentrations that exceed allowable pollutant discharge limits thereby adding the cost for both the disposal of the treated wastewater and the balance of plant costs of liquid electrolyte handling. Large electrode gaps and low surface area electrodes are also contributors to efficiency losses and increased energy consumption. Slow mass transport in the pores of the porous beds, non-optimized catalyst materials with poor reaction kinetics, high electrode overpotentials, and catalysts with low over potentials for side reactions (e.g. oxygen evolution) also contribute to lower performance and efficiency losses. The use of cell component materials which passivate quickly and increase cell resistivity and instabilities, contribute to efficiency losses. Operating conditions also contribute to efficiency losses. With high mass and ionic transfer losses, at nominal operating current densities, the voltages are too low such that incomplete destruction of organic contaminants occurs and an organic film blocks catalyst sites reducing performance and requiring the use of cell reversal techniques to clean the electrode surfaces.

For instance, published PCT application WO9901382 discloses an electrolytic cell method and apparatus for the decontamination of fluids. The system advantageously comprises means for adding one or more chemical substances into the fluid to be treated (e.g. an acid, carbon dioxide, an alkali, hydrogen peroxide, or a salt.) In another example, Andrade et al. in J. Haz. Mats. 153, 252-260 (2008) disclose the use of a divided electrolytic cell to treat model phenol wastewater. A supporting electrolyte of sulfuric acid was required.

To eliminate the requirement for supporting electrolyte addition, various methods have been developed that reduce the electrode gap in single compartment electrochemical cell configurations. For example, U.S. Pat. No. 6,328,875 discloses the use of a porous anode allowing wastewater to penetrate through the anode to flow through the capillary inter-electrode gaps. However, the energy consumption was still high when run without a supporting electrolyte. As with all single chamber electrochemical systems, hydrogen is simultaneously produced and wastewater constituents are reduced on the cathode, which consume much energy. Fouling of the cathode commonly occurs from these reaction products, decreasing the cell efficiency and leading to increased energy consumption. Another problem encountered in single chamber systems during oxidation is the production of intermediate compounds. These compounds are reduced at the cathode and are then reoxidized at the anode decreasing cell efficiency and increasing energy consumption.

An approach to eliminate the requirement for addition of a supporting electrolyte addition is to use a solid polymer electrolyte (SPE) in the electrolytic cell. SPE technology has been developed for other purposes including the production of hydrogen by water electrolysis and of energy using polymer electrolyte membrane fuel cells. For instance, in the system disclosed in WO03093535, dehalogenation of halogenated organic compounds and destruction of nitrates is conducted on the cathode by electrochemical reduction. In this configuration, the anode and cathode compartments are divided by an ion exchange membrane and an anolyte and halogen-containing catholyte are passed through their respective chambers. Although the system operated without supporting electrolytes, in order to operate at low current density (high cell efficiency), a supporting electrolyte was required in the anolyte and/or catholyte. Murphy et al. in Wat. Res. 26(4) 1992 443-451 used a SPE electrolytic cell to treat wastewaters with low or negligible supporting electrolyte content. The wastewater was re-circulated through both the anode and cathode. The energy consumption was very high however, and was attributed to low rates of phenol oxidation and side reactions, primarily oxygen evolution from water. J. H. Grimm et al. in J. Appl. Elect. 30, 293-302 (2000) used a SPE electrolytic cell to treat model phenol containing wastewater. The wastewater was pumped through the anode and cathode chambers in series. The energy consumption however was also high for phenol removal, which was attributed by the authors to the loss in current efficiency due to side reactions such as oxygen evolution. Further, A. Heyl et al. in J. Appl. Electrochem. (2006) 36:1281-1290 investigated a range of SPE electrolytic cell configurations at higher temperatures to de-chlorinate 2-chlorophenol model wastewater. In all cases, the wastewater was pumped across the membrane from either the cathode or anode to the opposite chamber through perforations in the membrane or by assisted electro-osmotic drag of treated membranes. The energy consumption was found to be impractically high for the untreated membrane, lower for the chemically treated membrane, and lowest for the perforated membrane. However, the best mineralization was obtained with anodic oxidation first followed by cathodic reduction with higher energy consumption. Still further, another approach for treating low conductivity wastewater without the use of supporting electrolytes was disclosed in WO2005095282. The system used a solid polymer electrolyte sandwiched between anode and cathode electrodes place in a single chamber of low conductivity wastewater. The energy consumption for pollutant mineralization of this setup was high due to the high voltages required.

Systems have also been developed in the art to reduce the cost of producing hydrogen by electrolysis by integrating electrolytic treatment of wastewater therewith. The electrolytic cells involved can use anolytes containing organic pollutants. For instance, Park et al. in J. Phys. Chem. C. 112(4) 885-889 (2008) used a single chamber cell to treat aqueous pollutants and produce hydrogen. As with all single chamber systems, a supporting electrolyte was required. The hydrogen generated was contained in a mixed product gas that required further treatment to recover usable hydrogen. Similar single chamber configurations were disclosed by T. Butt & H. Park in WEFTEC 2008 Conference Proceedings and by J. Jiang et al. in Environ. Sc. & Tech. 42(8), 3059 (2008). Divided cell configurations were disclosed for instance in WO2009045567 and by Navarro-Solis et al. in I J Hydrogen Energy 35 (2010) 10833-10841. The preceding systems all involved the use of additional supporting electrolytes. Systems without supporting electrolytes have also been disclosed for instance by F. Kargi in I. J. Hydrogen Energy 36 (2011) 3450-3456.

Systems using a solid polymer electrolyte based electrolytic cell have also been disclosed in the art to generate hydrogen and to treat wastewater. For instance, U.S. 65/333,919 discloses a method for electrolysis of an aqueous solution of an organic fuel. In this system, permeation of unreacted methanol to the cathode (fuel crossover) takes place and causing high cathode overpotentials and requiring the addition of a hydrogen gas cleaning operation. Further, E. O. Kilic et al. in Fuel Proc. Tech. 90 (2009) 158-163 disclose a system to treat formic and oxalic acid and generate hydrogen. However, the specific energy consumption was high due to the higher current densities required.

Notwithstanding the substantial developments in the art, there remains a continuing need for more efficient and cost effective methods for wastewater treatment. The present invention addresses this need while additionally providing other benefits as disclosed herein.

SUMMARY OF THE INVENTION

Methods and systems have been discovered for the energy efficient treatment of polluted wastewater using certain electrolytic cell designs and a combination of voltage and current density limitations. A lower current density results in better efficiency, and a lower voltage results in less unwanted side reaction (e.g. oxygen evolution). A higher flow rate results in lower energy consumption. Improved energy efficiency can be achieved while essentially removing all the pollutant.

The electrolytic cells employed comprise a solid polymer electrolyte electrolytic cell comprising an anode, a cathode, and a solid polymer membrane electrolyte separating the anode and the cathode. The anode comprises an anode catalyst layer, and the anode catalyst layer comprises an anode catalyst. In a like manner, the cathode comprises a cathode catalyst layer and the cathode catalyst layer comprises a cathode catalyst. The cathode in the electrolytic cell is liquid-electrolyte free. That is, the cathode comprises no liquid catholyte nor liquid supporting electrolyte.

The present method comprises providing at least a first and second solid polymer electrolyte electrolytic cell stack, supplying a flow of wastewater comprising a pollutant to the anode of each of the first and second electrolytic cell stacks at a flow rate and flow pressure, providing a voltage less than about 3 volts across each of the cells in the first and second electrolytic cell stacks wherein the anode is positive with respect to the cathode, operating each of the cells in the electrolytic cell stacks at an operating temperature and a current density less than about 20 mA/cm², and particularly less than about 10 mA/cm². This results in the pollutant being degraded and hydrogen gas being generated at the cathode. The generated hydrogen gas is exhausted from the cathode. The flow of wastewater can be supplied to the anode without an added supporting electrolyte, and the electrolytic cell can be operated over a wide range of wastewater temperatures. In particular here, either a stack component in each of the two stacks is different or an operating condition of the two stacks is different. In other words, at least one of a stack component in the first solid polymer electrolyte electrolytic cell stack and an operating condition of the first solid polymer electrolyte electrolytic cell stack is different from the stack component in the second solid polymer electrolyte electrolytic cell stack and the operating condition of the second solid polymer electrolyte electrolytic cell stack.

In embodiments in which a stack component differs between the first and second solid polymer electrolyte electrolytic cell stacks, the different stack component can be selected from the group consisting of an anode fluid delivery layer, the anode catalyst, the anode catalyst layer, an anode flow field plate, an anode filter layer, the solid polymer electrolyte membrane, and the number of cells in the stack.

In embodiments in which the different stack component is the anode catalyst layer, it can be at least one of the catalyst loading and catalyst active area in the anode catalyst layer of the first solid polymer electrolyte electrolytic cell stack that differs by more than about 5% from that of the catalyst loading and catalyst active area in the anode catalyst layer of the second solid polymer electrolyte electrolytic cell stack. In certain embodiments, it can be that the catalyst loading or catalyst active area in the anode catalyst layer of the first solid polymer electrolyte electrolytic cell stack differs by more than about 10% from that of the catalyst loading and catalyst active area in the anode catalyst layer of the second solid polymer electrolyte electrolytic cell stack.

In embodiments in which an operating condition differs between the first and second solid polymer electrolyte electrolytic cell stacks, the different operating condition can be selected from the group consisting of the flow rate of the wastewater, the flow pressure of the wastewater, the voltage, the operating temperature, and the current density. The different operating condition of the first solid polymer electrolyte electrolytic cell stack can differ by more than about 5% from that of the second solid polymer electrolyte electrolytic cell stack, and in certain embodiments the operating conditions can differ by more than about 10%.

In the methods and systems of the invention, the first and second solid polymer electrolyte electrolytic cell stacks may each comprise a single electrolytic cell. Alternatively, either or both of the first and second solid polymer electrolyte electrolytic cell stacks may comprise more than one electrolytic cell. Further, additional solid polymer electrolyte electrolytic cell stacks may be employed.

In one configuration of the invention, the first solid polymer electrolyte electrolytic cell stack is connected upstream and in series flow with the second solid polymer electrolyte electrolytic cell and the anode outlet from the first solid polymer electrolyte electrolytic cell stack is connected to the anode inlet of the second solid polymer electrolyte electrolytic cell stack second cell stack. In certain such embodiments, the first and second solid polymer electrolyte electrolytic cell stacks can share common end plates.

In another configuration of the invention, the first and second solid polymer electrolyte electrolytic cell stacks are connected in parallel flow with the second solid polymer electrolyte electrolytic cell and the supplied wastewater is divided between the anode inlets of the first and second solid polymer electrolyte electrolytic cell stacks.

In the methods and systems of the invention, one or more treatment units may be incorporated in the flow of wastewater. The treatment unit or units may be a filter, a degas unit, and a pH controller. Such treatment units may be incorporated at various locations throughout a system, including upstream or downstream of either the first or the second solid polymer electrolyte electrolytic cell stack.

The method is suitable for removing a variety of pollutants from wastewater, e.g. an organic or mixture of organics, inorganics such as ammonia or hydrogen sulfide, or mixtures of organics and inorganics. As demonstrated in the Examples, the method is suitable for removing an organic pollutant such as Acid Blue dye, phenol, acetaminophen, formic acid, ibuprofen, or a mixture of organic pollutants from Kraft pulp and paper mill effluent. Pollutants oxidized using the method include dissolved organics, biological oxygen demand (BOD), chemical oxygen demand (COD), total organic carbon (TOC), recalcitrant organics that remain after biological treatment processes, ammonia, dissolved gases (VOC light hydrocarbons and H₂S hydrogen sulfide), microorganisms, pathogens, and metal ions.

In a wastewater mixture of pollutants, the energy requirement and cell operating conditions for optimized decomposition and/or oxidation is not equal for all constituents. Also, different catalysts are able to accelerate oxidation and/or decomposition of specific constituents. Therefore, the decomposition and oxidation of wastewater pollutants can be optimized to lower treatment cost by combining stacks in series and parallel that operate at a combination of lower and higher voltage, current density, temperature, pressure, and flow rate.

Additionally, or alternatively, each of the cell stacks may comprise different component designs and materials, such as anode fluid delivery layers, anode filter layers, catalyst composition and loadings, flow field plate type and design, polymer electrolyte membranes, electrode active areas, and cell count. Again, each stack of electrolytic cells may be optimized for different types of contaminants that may be in the wastewater stream. For instance, as shown in the following Examples, certain anode catalysts have a higher affinity for oxidizing different contaminants. Therefore, in wastewater streams with several different contaminants, one skilled in the art will be able to determine the best anode catalyst (or membrane electrode assembly design) for a particular type of contaminant, and use them in separate groups of segmented electrolytic cells or stacks of electrolytic cells for improved contaminant removal from the wastewater stream.

Embodiments of the system may comprise multiple electrolytic cells in stacks, in either series and/or parallel flow arrangements. For example, wastewater can be split and supplied to multiple stacks of cells and the flows combined thereafter at the cell or stack outlets. Each of the stacks of electrolytic cells may be different, for example, operating at different operating conditions and/or comprising different components. This is particularly useful for wastewaters that comprise appreciable concentrations of oxidizable constituents such as ammonia, hydrogen sulfide, metals and inorganics or constituents that decompose at low electrolytic potentials, decompose with heat, and with different catalysts. By optimizing the design of each stack for particular types of contaminants that may be in the wastewater stream, improved contaminant removal, energy efficiency, operating costs and cell lifetime may be achieved. As shown in the Examples below, cell (and membrane electrode assembly) designs and operating conditions have a large effect on removal of different contaminants.

Pollutant specific decomposition and oxidation catalysts may be desirably incorporated into the anode flow field plate, an anode filter layer, the anode fluid diffusion layer, or anode catalyst layer. These provide for the decomposition and/or oxidation of the pollutants at lower voltage, higher flow rates and lower energy consumption. Pollutant specific decomposition catalysts may be desirably incorporated into the catalyst layer to provide for faster rate of pollutant removal. This may allow higher flow rates and lower cell active area required to reduce energy consumption. The anode catalyst layer may alternatively only contain catalyst particles that speed up decomposition of a specific pollutant, e.g. MnO₂ decomposes H₂O₂, but does not react with the other pollutants in the wastewater. It may be beneficial to remove the resulting products quickly in this manner before oxidizing the remaining organics and inorganics. For example, if the products are gases and/or solids, the wastewater can be advantageously de-gassed or filtered in an intermediate step to prevent them from interfering with downstream processes.

The total energy required to remove pollutants from a mixed wastewater can be reduced by configuring the system to remove the easily decomposed or oxidized pollutants at lower energy first. The wastewater may pass through a series of stacks, each with a catalyst layer that targets one or more pollutants to decompose and/or oxidize them at low energy. For wastewater with rapid oxidation and decomposition of pollutants, the flow rate may be increased and the total cell area or number of cells reduced to lower cost and system footprint.

For pollutants that oxidize and/or decompose into gases, one or more degas units or methods may be employed in the system to remove resulting product gases. Dissolved gases (e.g. CO₂, O₂) may need to be removed due to corrosion and/or undesirable reactions in downstream equipment and processes. For example, in water with low concentrations of minerals, carbon dioxide forms carbonic acid which is corrosive. Degasification methods include heating (e.g. deaerating heaters), reducing pressure (e.g. vacuum deaerators), membrane processes (e.g. membrane contactors), air stripping, substitution with inert gas (e.g. bubbling with argon), vigorous agitation, contact with catalytic resins, and freeze-thaw cycling. For dissolved oxygen, chemical oxygen scavengers may also be added (e g ammonium sulfite). For dissolved carbon dioxide additional methods of removal include contact with limestone and/or magnesium oxide (to form carbonates and bicarbonates), chemical reaction with a solution of sodium carbonate to form sodium bicarbonate, and carbonic acid neutralization by controlling the pH between 7.5 and 8.5.

The following are operating conditions that may be adjusted individually or in combination from stack to stack in a system to provide for lower energy consumption:

-   -   the voltage may be reduced in one or more stacks for easily         decomposed or oxidized species to lower the energy required in         one or more stacks,     -   the pressure may be increased in one or more stacks to keep         gases dissolved,     -   the flow rate may be increased in one or more stacks,     -   the temperature may be increased in one or more stacks to         increase reaction kinetics,     -   the current density may be reduced in one or more stacks.

For pollutants that oxidize and/or decompose into gases, one or more degas units or methods may be employed in the system to remove resulting product gases. The method can comprise a degas step and/or an intermediate step between stacks, and/or a post-degas step.

The method can comprise a pre-filtration step and/or intermediate step between stacks, and/or a post-filtration step.

The method can comprise a combination of stacks in which certain cells comprise anode filter layers and others do not. The anode filter layers may be either electrically conductive or non-conductive. The anode and cathode flow field plates may also be electrically conductive or non conductive or a combination of both. The anode catalyst layers may have different compositions of catalysts and different concentrations of ionomer.

The method can comprise a pre-pH adjustment step and/or an intermediate step between stacks, and/or a post-pH adjustment step. This is advantageous for removing pollutants e.g. increase pH to precipitate metals or decrease pH to precipitate silica and for reducing corrosiveness of wastewaters with acids to permit the use of less corrosion resistant cell components and thereby reducing stack cost.

The method can additionally comprise a post treatment step for removing free chlorine selected from the group consisting of: reducing electrochemically, adsorbing, decomposing by contacting a transition metal, reacting with a salt, reacting with a chemical reducing agent, reacting with organic matter, decomposing by contacting a redox filter, decomposing by light exposure, and decomposing by heating.

Further, the method can comprise a step for preventing formation of chlorine selected from the group consisting of: controlling the pH of the wastewater to be greater than about 2, increasing the ionomer concentration at the anode fluid delivery layer, increasing the ionomer concentration at the anode catalyst layer, and incorporating materials that are known to catalyze the decomposition of free chlorine into the anode. The latter materials can include transition elements such as iron, copper, manganese, cobalt and nickel, Raney metals of copper, nickel and cobalt, their oxides and spinels and can be mixed into the anode catalyst layer. Alternatively, such materials can be applied as coatings to the anode fluid delivery layers and/or anode plates to effect decomposition of free chlorine.

Further still, the method can additionally comprise a cleaning step selected from the group consisting of: purging the anode and cathode with a cleaning solution, flowing solid particles through anode flow-field, in-situ backwashing, oxygen scouring, chemical cleaning, ultrasonic cleaning, gas purging, liquid purging, potentiostatic cleaning, flowing water, and generating chlorine and intermediates of oxygen at higher anodic voltages.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 shows a schematic of one embodiment of the inventive system and was used to perform the laboratory scale wastewater treatment in the Examples.

FIG. 2 shows a schematic of the solid polymer electrolyte cell used in the system of FIG. 1.

FIG. 3 shows a schematic of an alternative embodiment of an electrochemical cell suitable for use in the inventive system.

FIG. 4 shows a schematic of an embodiment of the inventive system having more than one electrochemical stack.

FIG. 5 shows a schematic of an embodiment of an electrochemical cell stack in cross section suitable for use in the inventive system, wherein the stack comprises segmented cell groups.

FIG. 6 shows a schematic of another embodiment of the inventive system having more than one electrochemical stack.

FIG. 7 shows a schematic of yet another embodiment of the inventive system having more than one electrochemical stack.

FIG. 8 is a qualitative prior art illustration showing how the change in original compound concentration can differ from that of the COD over the course of oxidation for refractory organic compounds such as phenol.

FIG. 9 compares the average actual hydrogen generated from a number of tests performed at several different currents on phenol contaminated wastewater to ideal or perfect hydrogen generation.

DETAILED DESCRIPTION

Certain terminology is used in the present description and is intended to be interpreted according to the definitions provided below. In addition, terms such as “a” and “comprises” are to be taken as open-ended. Further, all US patent publications and other references cited herein are intended to be incorporated by reference in their entirety.

Herein, SPE stands for solid polymer electrolyte and can be any suitable ion conducting ionomer, such as Nafion®. A SPE electrolytic cell is thus a cell comprising a SPE as the electrolyte to which electrical energy is supplied to effect a desired electrochemical reaction (with a positive voltage being applied to the anode of the cell).

Herein, unless otherwise specified, when referring to a numerical value the term “about” is intended to be construed as including a range of values within plus or minus 10% of the value being referred to.

An electrode in the cell is “liquid-electrolyte free” means that no significant ion containing liquid is deliberately provided to the electrode, such as is done in certain systems in the prior art. However, it is not intended at the cathode for instance to exclude minor amounts of wastewater which may cross over through a solid polymer electrolyte.

An “electrolytic cell stack” refers to a series stack of electrolytic cells comprising one or more electrolytic cells.

A “stack component” refers to any of the components making up a solid polymer electrolyte electrolytic cell stack of the invention. It includes but is not limited to an anode fluid delivery layer, the anode catalyst, an anode filter layer, the anode catalyst layer, an anode flow field plate, the solid polymer electrolyte membrane, and the number of cells. Further, it includes any special sublayers employed, the cathode catalyst, the cathode catalyst layer, a cathode gas diffusion layer, and a cathode flow field plate.

An “operating condition” refers to any of the variable operating conditions employed in the operation of a solid polymer electrolyte electrolytic cell stack of the invention. It includes but is not limited to the flow rate of the wastewater, the flow pressure of the wastewater, the voltage, the operating temperature, and the current density.

The energy efficient system of the invention employs a simple, compact electrolytic cell architecture to minimize ionic, ohmic and mass transport resistances, and is characterized by a reduced operating voltage and current density, low-cost components, a chemically stable, non-liquid electrolyte membrane, and low-cost, durable and high performance electrode and catalyst designs. Recovery of high purity, by-product hydrogen is possible for enhanced efficiency.

An exemplary system is shown in the schematic of FIG. 1. System 100 comprises SPE electrolytic cell 101 for the direct electrochemical treatment of organic contaminated wastewater. A controlled flow of wastewater 102 is supplied to anode inlet 11 of cell 101 by some suitable delivery means, e.g. peristaltic pump 103. After sufficient treatment/transit time in cell 101, the treated wastewater exits at anode outlet 12 and as shown here is delivered to treated effluent tank 104 where entrained or by-product gases generated during treatment (e.g. carbon dioxide, nitrogen, oxygen) are allowed to vent to atmosphere. For monitoring and flow control purposes, pressure gauge 105, valve 106, and flowmeter 107 are provided in the anode outlet line.

Electrical energy is provided to cell 101 by DC power supply 108 and the temperature of cell 101 is monitored and controlled by temperature controller 109. Hydrogen is generated at the cathode of cell 101 as a result of the electrochemical treatment and is exhausted at cathode outlet 13. As shown in FIG. 1, the relatively pure hydrogen may be collected and stored in storage container 110 for later use in the generation of electricity or as a fuel or chemical reactant.

FIG. 2 shows a detailed schematic of solid polymer electrolyte electrolytic cell 101. Cell 101 comprises SPE membrane electrolyte 6. The cell anode comprises anode catalyst layer 8 and anode fluid delivery layer 9. The cell cathode comprises cathode catalyst layer 5 and cathode gas diffusion layer 4. Anode flow field plate 10 is provided adjacent anode fluid delivery layer 9. Anode flow field plate 10 comprises flow field channel/s 10 a which are fluidly connected to anode inlet 11 and anode outlet 12. Wastewater 101 is delivered uniformly to and from anode fluid delivery layer 9 by directing it through flow field channel/s 10 a. Cathode flow field plate 3 is provided adjacent cathode gas diffusion layer 4. Cathode flow field plate 3 comprises flow field channel/s 3 a which are fluidly connected to cathode outlet 13. Since no catholyte nor other liquid nor fluid is supplied to the cathode, a cathode inlet is not required. Hydrogen gas generated during the electrochemical treatment of wastewater 101 however is collected from the cathode and directed out of the cell by way of flow field channel/s 3 a. Leads 2 are provided at the cell electrodes in order to make electrical connections to power supply 108. Mechanical support is provided to the components in cell 101 by way of end plates 1 which clamp the cell together. Sealing is provided to the cell by seals 7. A drain port (not shown in FIG. 2) may be incorporated at the cathode for purposes of cleaning and/or purging any water crossover accumulation. Finally, FIG. 2 shows heating elements 14 which may be used to control the cell temperature during operation.

Cell 101 may optionally include an anode filter layer (not shown) which is incorporated between anode flow field plate 10 and anode fluid delivery layer 9. During assembly, such a filter layer may be applied to anode fluid delivery layer 9, or applied to anode flow field plate 10, or applied as a discrete component to be clamped in place by end plates 1. Such filter layers may be provided to prevent particulate and suspended solids contaminants from entering anode fluid delivery layer 9. The average pore size of such a filter layer is chosen to remove the particles expected to be found in the wastewater. For example, to filter particles of 50 micrometer size or greater, the average pore size would be less than 50 um. Further, the average pore size of a filter layer is preferably smaller than the average pore size of the adjacent anode fluid delivery layer 9. In situations where the pore size throughout anode fluid delivery layer 9 is not uniform, the average pore size of a filter layer is preferably less than that of the flow field side of anode fluid delivery layer 9. This configuration provides for removal of contaminants while still allowing filtered wastewater to access the anode catalyst layer for treatment. Cake formation on the surface of a filter layer is promoted and this prevents clogging of both the filter layer and anode fluid delivery layer 9. Also oxidation products can readily be removed from the catalyst layer.

Unexpectedly high energy efficiency can be obtained from system 100 and can result from appropriate limitations to the voltage and current density applied to the cell and by adoption of some of the designs and components used in advanced SPE fuel cells for the generation of electricity. In particular, the voltage applied across electrolytic cell 101 (or across individual cells if more than one is employed in a system) should be less than about 3 volts. The current density is limited to below about 20 mA per cm² of electrode area. And as discussed further below, certain catalyst choices, catalyst layer constructions, fluid delivery layer and gas diffusion layer constructions can benefit operating efficiency.

The reasons for the improved efficiency of the instant invention are not completely understood. However, without being bound by theory, several mechanisms may be involved at the anode for the mineralization of organic pollutants. Oxygen for the “electrochemical incineration” of organic pollutants in the wastewater is obtained from water from an oxygen evolution reaction. Adsorbed hydroxyl radical and oxygen radical species generated on the surface of the anodic catalyst can mineralize organic pollutants present. In addition, for certain n-type semiconductor oxide catalyst, anionic (oxygen) vacancies can preferentially react with water and generate OH* radicals. Oxidation via intermediates of oxygen evolution/hydroxyl radicals at anodic potentials in the region of water discharge can mineralize or partially oxidize organic pollutants. Direct oxidation of ammonia to nitrogen may occur. Further, indirect electrochemical oxidation may take place by inorganic oxidants generated by anodic oxidation of sulfate, carbonate, or phosphate ions in the wastewater. Decomposition of pollutants may occur. Direct electrolysis of pollutants may occur. And further still, there may be indirect electrochemical oxidation by redox reagents electrochemically generated from a mediator present in the wastewater.

The chemical reactions involved at the anode can include:

Direct electrolysis of organic compound R by electron transfer: R→P+e ⁻

For the mineralization of organic compounds, R, through oxygen transfer from water and evolved oxygen:

${R + {\frac{n}{2}H_{2}O}}->{{{mineralization}\mspace{14mu}{{products}\mspace{14mu}\left\lbrack {{{CO}_{2} + {salts}},{{etc}.}} \right\rbrack}} + {nH}^{+} + {ne}^{-}}$   2H₂O− > O₂ + 4H⁺ + 4e⁻ ${R + {\frac{n}{4}O_{2}}}->{{{mineralization}\mspace{14mu}{{products}\mspace{14mu}\left\lbrack {{{CO}_{2} + {salts}},{{etc}.}} \right\rbrack}} + {{nH}^{+}{ne}^{-}}}$

For hydroxyl and oxygen radicals, and intermediates of O₂ evolution on a catalyst surface: H₂O→OH*_(ads)+H⁺ +e ⁻ (h ⁺)_(vac)+H₂O→(OH*)_(ads)+H⁺ +e ⁻ R+[OH*radicals/O*species/intermediates]_(ads)→mineralization products [CO₂+salts, etc.]+nH⁺ +ne ⁻

For the oxidation of ammonia 4NH₃+3O₂→2N₂+6H₂O NH₃/NH₄+OH*→N₂+H₂O+H⁺ +e ⁻ and if the wastewater is alkaline, removal via free chlorine HOCl+⅔NH₃→⅓N₂+H₂O+H⁺+Cl⁻ NH₃/NH₄+HOCl/OCl⁻→N₂+H₂O+H⁺+Cl⁻ For the formation of inorganic oxidants, e.g.: 2CO₃ ²⁻→C₂O₆ ²⁻+2e ⁻ 2PO₄ ³⁻→P₂O₈ ⁴⁻+2e ⁻ 2HSO₄ ⁻→S₂O₈ ²⁻+2H⁺+2e ⁻

For the generation of oxidants in-situ, e.g. NaCl in wastewater:

2Cl⁻− > Cl₂ + 2e⁻ ${{\frac{1}{2}{Cl}_{2}} + {H_{2}O}}->{{HOCl} + H^{+} + {Cl}^{-}}$ HOCl− > H⁺ + OCl⁻ For  H₂S: H₂S− > S^(∘) + 2H⁺2e⁻

And if the wastewater is alkaline, via electrochemical decomposition (see “A Modified Electrochemical Process for the Decomposition of Hydrogen Sulfide in an Aqueous Alkaline Solution”, Z. Mao, A. Anani, Ralph E. White, S. Srinivasan & A. J. Appleby. Journal of the Electrochemical Society, 1991, pages 1299-1303.) A pH control apparatus may be employed to facilitate alkaline decomposition

And for metal ions [e.g. transition metal ions such as iron, manganese]:

oxidization via hydroxyl radicals and oxygen

oxidation via hydroxyl radicals, e.g. Mn+OH*→Mn⁻¹+OH⁻

or oxidation with oxygen, e.g. 2Fe²⁺+½O₂+5H₂O→2Fe(OH)₃↓+4H⁺ Mn²⁺+½O₂+H₂O→MnO₂↓+2H⁺

For such purposes, oxygen generating electrocatalysts may desirably be incorporated into a catalyst layer deposited on a fluid diffusion layer. Further, the residence time of wastewater in contact with the catalyst layer may be increased to complete oxidation, and a microfilter may be employed in the system to remove resulting metal precipitates.

And catalytic decomposition: H₂O₂→H₂O+½O₂

Pollutant specific decomposition and oxidation catalysts may be desirably incorporated into the anode flow field plate, an anode filter layer, the anode fluid diffusion layer, and anode catalyst layer. These can provide for the decomposition and/or oxidation of the pollutants at lower voltage, higher flow rates and lower energy consumption.

Pollutant specific decomposition catalysts may be desirably incorporated into the anode catalyst layer to provide for faster rate of pollutant removal at higher voltages. This may allow higher flow rates and lower required cell active areas so that smaller stacks at lower cost may be employed.

For pollutants that oxidize and/or decompose into gases, one or more degas units or methods may be employed in the system to remove resulting product gases.

Meanwhile at the cathode, hydrogen evolution occurs as:

${{nH}^{+} + {ne}^{-}}->{\frac{n}{2}{H_{2}(g)}}$

Kinetic effects generally are believed to dominate at the low current densities involved in the present method, and thus the catalysts used may have a great effect. A high active surface area may allow more OH radicals to be available, the electron and proton transfer media present (e.g. conductive particles and ionomer) enhance charge transfer, and additional particles may also contribute to generate local oxygen (e.g. high surface area graphite particles). The use of advanced fuel cell components may also assist in improved mass transfer and current collection and local mixing of fluids at the catalyst surfaces if there is not excessive oxygen generation at the anode.

In the present invention, there may be a preferred amount of oxygen produced where too little means not enough is present for the pollutant removal related reactions to take place at a reasonable rate and yet where too much oxygen production is parasitic and the current density shoots up while the rate of contaminant removal remains the same. In the list of preceding anode reactions, the mineralization of organic compound reactions is frequently cited in the literature. However, the reactions for hydroxyl and oxygen radicals, and intermediates of O₂ evolution on a catalyst surface may be of importance. A small amount of locally generated oxygen may occur on alternative particles without compromising catalytic sites for OH radicals. In effect, this may result in increased reaction kinetics, and the same organic pollutant removal rate might be achieved at lower applied voltage and current densities. For electrodes in the prior art, in order to obtain a decent level of OH radicals, the applied voltage needs to be increased thereby driving the cells into a substantial range for oxygen production that may then compete with radical production sites. That is, higher voltages and current densities may be needed in the prior art to get an equal amount of OH radicals.

Regardless, unexpected improved energy efficiency has been obtained when appropriately limiting the applied voltage and current density as mentioned previously and also by using certain electrolytic cell designs and components. Energy efficiency can be further improved by using a combination of electrolytic stacks at limited applied voltage and current along with faster wastewater flow rates. SPE membrane electrolyte 6 is a suitable proton conducting solid polymer electrolyte and is preferably a thin, extended life material choice to increase efficiency (e.g. sulfonated tetrafluoroethylene based fluoropolymer-copolymer such as Nafion® in a thickness less than about 30 micrometers). However, for durability and/or high temperature service, membrane electrolyte thickness may desirably be increased to between 50 and 100 micrometers (e.g. by laminating thinner membranes together or using thicker membranes).

With regards to the anode catalyst, platinum, tin oxide, antimony tin oxide, manganese oxide and mixtures thereof have been used successfully in the Examples. In the case of antimony tin oxide, heat treatment to improve its electrical conductivity or doping, for instance with Nb, may be considered to improve durability. Manganese oxide can be considered for purposes of decomposing any hydrogen peroxide which may be formed at the anode. Other n- and p-type semiconductor oxides, perovskite-like oxide classes, and amorphous or nanocrystalline transition metal oxides (e.g. MoO₂) may also be considered as anode catalysts. Further, spinels of cobalt and nickel, and high surface area nickel oxides may also be considered. As is known in the art, use of supported catalysts (e.g. Pt dispersed on carbon or antimony tin oxide on high surface area graphite or Nb particles) can improve the dispersion of the catalytic materials and thus utilization and also the interaction between certain catalysts and supports can enhance catalytic activity and durability. Generally dopants can be employed to improve electrical conductivity (e.g. antimony doped SnO₂, chlorine and fluorine doped SnO₂) or to improve durability and stability at elevated voltages (e.g. cobalt, nickel, palladium, niobium, tantalum, platinum, palladium, iridium, ruthenium, vanadium, rhenium), and mixtures of such dopants to improve both electrical conductivity and stability/durability (e.g. SnO₂ doped with Nb and a dopant selected from the group Sb, Fe, F, Pt and Ni). Other possible dopants include Mo, Cr, Bi, and W.

Pollutant specific decomposition and oxidation catalysts may be desirably incorporated into the anode flow field plate, an anode filter layer, the fluid diffusion layer and/or the anode catalyst layer. These provide for the decomposition and/or oxidation of the pollutants at lower voltage, higher flow rates and lower energy consumption. Pollutant specific decomposition catalysts may be desirably incorporated into the catalyst layer to provide for faster rate of pollutant removal at higher voltages. This may allow higher flow rates and lower cell active area required so smaller stacks and lower capital cost.

Depending on the composition of the wastewater, the decomposition and pollutant specific catalysts may be MnO₂, PbO₂, Fe₂O₃, metal oxides, mixed metal oxides, doped metal oxides, metal oxides and mixed metal oxides and doped metal oxides supported on zeolites, aluminum oxides, calcium oxides and potassium oxides; metal sulfides, bimetallic catalysts such as Ni—Pt and Pt—Sn, cobalt compounds, transition metals and alloys, Pt, Au, Ru, Ir, Pd, Au and their compounds, noble metals supported on metal oxides such as alumina, silica, & ceria, transition metal-exchanged zeolites, perovskites, metalloporphyrins, silver compounds, activated carbon, Cu, CeO₂ promoted cobalt spinel catalysts, carbides, nitrides and functionalized silica catalysts.

The composition of the catalyst layer may also include higher concentrations of ionomer as it acts as an acid catalyst.

The selected catalyst materials are catalytic at lower voltages for the organic contaminants (i.e. have a lower overpotential) so the applied voltage required is lower and consequently, the current density is lower. Such catalyst materials have a high overpotential for water electrolysis, so that the generation of oxygen can be controlled at the operating voltage thereby reducing the current density associated with this reaction.

Other considerations in the selection of anode catalyst include use of nanoparticles, nanostructured and/or mesoporous materials to obtain high surface areas. Supported catalysts may be employed using supports of graphite. If stability of graphite at elevated anodic voltages is an issue, stable, conductive particles including carbides, nitrides, borides, corrosion resistant metals, alloys, and metal oxides (e.g. Nb, Nb₂O₅, ZnO, NbC and/or mixtures thereof) can be employed. Additives can include perovskite-based metal oxides that exhibit mixed electronic and ionic conductivity.

Anode catalyst layer 8 generally comprises particles to improve electron conduction, ionomer (e.g. similar to that used in the membrane electrolyte) for ion conduction and to serve as a binder, and material to control the wetting characteristics (e.g. dispersed PTFE). Pore size and overall porosity can be engineered to some extent by choice of particle size and agglomerate size (which can be modified for instance by controlling the high shear mixing rate during preparation of a catalyst ink or slurry used to make the catalyst layer). The pore characteristics of the anode catalyst layer, the surface chemistry and surface area can be important with regards to the mass transport of wastewater to the catalyst and the removal of product gas such as carbon dioxide. Preferably, the pore structure and hydrophobic surfaces of the anode catalyst layer facilitate bubble detachment so that gas blanketing and/or pore blockage does not occur. A graded particle size and pore size distribution can be employed in catalyst layer 8 to allow deeper penetration of wastewater and greater catalyst surface area utilization.

Anode fluid delivery layer 9 is provided to readily deliver fluids to and from anode catalyst layer 8 in a uniform manner. In addition, it provides electrical contact and mechanical support thereto. Carbon fibre paper, foams, and other materials commonly employed in SPE fuel cell embodiments may be contemplated here as substrates. And materials for electrical conduction and wettability may be added thereto. As with anode catalyst layer 8, the pore size distribution and bulk porosity of anode fluid delivery layer 9 is carefully controlled as it can be important with regards to carbon dioxide bubbles formed (effecting size and mixing) and their effect on mass transport. Sublayers (not shown in FIG. 2) commonly used in fuel cell embodiments may be incorporated in anode fluid delivery layer 9 and located adjacent to anode catalyst layer 8 in order to improve contact to the latter and to provide an asymmetric pore size distribution across layer 9 (e.g. to provide larger pores on the side adjacent anode flow field plate 10 which may act as a pre-filter preventing suspended solids from blocking catalyst sites).

If elevated anode potentials are involved, dissolution of materials such as carbon fiber paper may occur. In such cases, more stable media can be employed including metal coated (e.g. nickel coated) carbon fiber paper or woven cloth, metal screen/gauze/cloth (particularly with 2 or more ply screens with different mesh sizes and the smaller closest to membrane, with weave patterns to promote in-plane water permeability, flattened and diffusion bonded or spot welded together), sintered metal screen/gauze/cloth (again with 2 or more ply screens to improve current distribution and flattened), expanded metal foil/film/membrane (with 1 or more plies and flattened), sintered metal fiber and powder media (again with 1 or more plies and flattened, having asymmetric pore size and with the smaller pore diameter located closest to membrane, and having high in-plane water permeability), flattened photo-etched media, chemically etched media, micro-perforated plate, or combinations thereof. The preceding materials are electrically conductive and can be corrosion resistant types [stainless steel, inconel, monel, titanium, alloys, valve metals] or have corrosion resistant coatings applied thereto [e.g. carbides, nitrides, borides, noble & valve metals & metal alloys, metal oxides]. Conductive coatings may be applied to the surfaces contacting the catalyst coated membrane if the corrosion resistant materials employed form passive layer. Sublayers can be applied incorporating corrosion resistant and electrically conductive particles [e.g. carbides, nitrides, borides, noble & valve metals & metal alloys, metal oxides]. For monopolar designs, high in-plane conductivity is desirable, suggesting use of corrosion resistant, conductive, materials and coatings therefor.

The cathode catalyst can be selected from the group of conventional catalysts commonly used for hydrogen evolution, including platinum or supported platinum (e.g. carbon supported platinum), palladium, palladium alloys, supported Pd/C, nickel & oxides thereof, rhodium (e.g. metals where significant coverage by H₂ species is possible), molybdenum disulfide, perovskite-based metal oxides that exhibit mixed electronic and ionic conductivity, amorphous or nanocrystalline transition metal oxides, and high surface area Raney metals and metal blacks. In addition, manganese oxide, graphite, and carbon may also be employed at the cathode. Again, manganese oxide may be beneficial to decompose any hydrogen peroxide present. Along with cathode catalyst, cathode catalyst layer 5 also generally can comprise particles to improve electron conduction, ionomer for ion conduction and to serve as a binder, and material to control the wetting characteristics. Cathode catalyst layer 5 can be prepared by coating onto cathode gas diffusion layer and sintering at an appropriate temperature (e.g. about 150° C. or 370° C. respectively depending on whether ionomer or PTFE is employed). Conductive particles in layer 5 can desirably be mixed to provide a size distribution that optimizes current distribution and porosity for hydrogen recovery. If erosion is an issue, PTFE and/or other stable binders in catalyst layer 5 can be employed for improved erosion/wear resistance.

Cathode gas diffusion layer 4 is provided to readily deliver gases to and from cathode catalyst layer 5 in a uniform manner Layer 4 is desirably designed to repel wastewater which may cross-over from the anode side through the membrane electrolyte, while still permitting ready removal of generated hydrogen gas. Thus, a hydrophobic construction may be employed, for instance a teflonated stainless steel mesh substrate. Further, use of a hydrophobic sublayer with a small pore structure adjacent cathode catalyst layer 5 may also serve to prevent wastewater cross-over from entering the rest of the cathode. In turn, this can reduce or eliminate parasitic reactions and contamination at the cathode and thereby help keep the current density low. In general, materials similar to those employed in anode fluid delivery layer 9 may be considered. For monopolar designs, high in-plane conductivity is desirable, suggesting use of corrosion resistant and hydrogen resistant, conductive, materials and coatings therefor (e.g. nickel, palladium alloys, titanium nitride, etc.).

The flow field channels 3 a, 10 a in the cathode and anode flow field plates 3, 10 can have numerous configurations, including single serpentine, interdigitated, and/or multiple linear designs, and the cross-sections can have various shapes. Designs for gravity assist may be employed. Accommodating the hydrogen generated at the cathode is relatively straightforward and one end of the cathode flow field may be dead-ended. At the anode, channel design preferably maximizes residence and encourages uniform mixing of the liquids and generated gases. It can be useful to provide for turbulence to promote the mixing of gas and liquid and to prevent bubble coalescence and large plugs of gas from forming. This may be accomplished by providing static means for in-line mixing in the channels, e.g. spiral tape, twisted tape, or helical static mixing elements in various locations within flow field channels 10 a. Such mixing can serve various purposes including reducing a concentration overvoltage at anode, eliminating radial gradients in temperature, velocity and material composition, and improving mass transport of the wastewater allowing larger channels and higher wastewater flows to be used without any loss to performance. Appropriate mixing components would continuously mix the wastewater and direct the wastewater flow to the outer perimeter so that pollutants are efficiently delivered to the catalyst layer and gas bubbles are contacted with the porous plate surfaces for removal.

FIGS. 1 and 2 depict one possible embodiment of the system and electrolytic cell and versions of this were used in the Examples to follow. However, many other variations are possible and include a monopolar cell design comprising non-conducting plastic plates with conductive film on landings for current collection or with a metal substrate used in the anode fluid delivery layer for current collector. Other monopolar and bipolar variations may be contemplated including bipolar pairs within a monopolar stack. Plate materials in such cases can be varied. In monopolar designs, plates can be electrically insulating and made of plastic, composite (e.g. glass fiber reinforced plastic), ceramic, or metals coated with insulating, corrosion resistant coatings. In bipolar designs, plates are electrically conductive and can be made of composites (carbon plastic, fiber reinforced where fibers are conductive metals, carbides, nitrides, etc.), metals, alloys, and substrates comprising appropriate coatings (similar to those of anode delivery layer 9 on the anode side and gas diffusion layer 4 on the cathode side). In a monopolar stack comprising bipolar pairs, an electrically conductive cathode plate can be employed in between two electrically insulating anode plates.

Dissolved gases (e.g. CO₂, O₂) may need to be removed due to corrosion and/or undesirable reactions in downstream equipment and processes. For example, in water with low concentrations of minerals, carbon dioxide forms carbonic acid which is corrosive. Degasification methods include heating (e.g. deaerating heaters), reducing pressure (e.g. vacuum deaerators), membrane processes (e.g. membrane contactors), air stripping, substitution with inert gas (e.g. bubbling with argon), vigorous agitation, contact with catalytic resins, and freeze-thaw cycling. For dissolved oxygen, chemical oxygen scavengers may also be added (e g ammonium sulfite). For dissolved carbon dioxide additional methods of removal include contact with limestone and/or magnesium oxide (to form carbonates and bicarbonates), chemical reaction with a solution of sodium carbonate to form sodium bicarbonate, and carbonic acid neutralization by controlling the pH between 7.5 and 8.5.

Also possible are designs employing a porous anode plate, e.g. porous graphite or porous metal plates with small pores for degassing the wastewater. In such a design, the channel surfaces can be made hydrophobic to prevent water ingress with the maximum pore size dependent on contact angle of plate surface and operating pressure of the wastewater flow. FIG. 3 shows a schematic of such an alternative embodiment 111 based on a porous anode plate option. (In FIG. 3, like numerals have been used to indicate components similar to those shown in FIGS. 1 and 2.) Here, the electrolytic cell comprises porous anode plate 15 and gas collection manifold 16. A vacuum assist at the anode outlet is also provided by vacuum pump 17 to assist in the removal of gases. Other options include the use of a 2-stage system, instead of a single stage, in which two electrolytic cells are employed in series with the anode outlet from one being connected to the anode inlet of the other, and in which generated hydrogen is collected from both cathodes.

The energy efficient benefits of the invention are obtained by limiting the current density and the voltage applied per electrolytic cell in the system. Other operating conditions are fairly flexible. Any operating temperature between the freezing point and boiling point of the wastewater may be considered (e.g. from about 3 to 95° C.) although temperatures modestly elevated above ambient may be useful in increasing reaction rates (e.g. from about 25 to 50° C.). Wastewater may typically be supplied at pressures from about 0 to 30 psi. The transit time or residence time of the wastewater is selected in order to ensure adequate removal of pollutants from the wastewater.

Depending on what is specifically in the wastewater, certain modifications can be considered. For instance, if the wastewater contains acid, base, alkali and/or other ionic species that make it conductive, ionomer may not be required in the catalyst layer and an alternative binder may be employed (e.g. PTFE). If high chloride ion levels are present in the wastewater, it may react at anode electrocatalytic sites to produce free chlorine (defined as dissolved Cl₂ gas, hypochlorous acid HOCl and/or hypochlorite ion OCl⁻ in equilibrium together and whose concentrations are a function of pH). Here, pH may be controlled to prevent dissolved Cl₂ gas (pH>2). And divalent ions can be added to the wastewater to increase the concentration therein (such as sulphate SO₄ ²⁻ and/or sulphate salts such as NaSO₄). Such divalent ions preferentially adsorb onto the electrode, catalyze oxygen formation, and inhibit the oxidation of chloride ions. Further, transition elements such as iron, copper, manganese, cobalt and nickel, Raney metals of copper, nickel and cobalt, their oxides and spinels can be mixed into the catalyst layer that are known to catalyze the decomposition of free chlorine. Such materials can be applied as coatings to the anode fluid delivery layers and/or anode plates to effect decomposition of free chlorine. Further, a post treatment step may be employed to remove free chlorine, including: electrochemical reduction, adsorption by granular activated carbon or kaolinite clay, decomposition by contacting transition metals (especially copper, iron, nickel and cobalt and/or their oxides and spinels such as substituted cobalt oxide spinels), reacting with salts such as ammonium acetate, ammonium carbonate, ammonium nitrate, ammonium oxalate, and ammonium phosphate, reacting with chemical reducing agents such as sodium metabisulfite, reacting with organic matter such as glycerol, decomposition by contacting redox filters such as copper/zinc alloys, decomposition by light exposure (especially UV), and decomposition by heating the solution. Further still, the ionomer concentration at the anode fluid delivery layer or catalyst layer may be increased to block chloride ions from catalytic reaction sites.

In certain cases during operation, species can undesirably migrate into regions of the electrolytic cell. For instance, if the wastewater contains high levels of metallic ions that are not all oxidized, a portion can diffuse into the membrane. This problem may be addressed by performing an in-situ ion exchange cleaning procedure, or alternatively a pre-treatment step may be employed to remove or reduce these via chemical coagulation-flotation/filter/clarifier, electro-coagulation & flotation/filter/clarifier, lime softening, chemical precipitation, and so on. Further, one or more of the following may be performed to reduce fouling and cleaning requirements: removal of suspended solids, particulate matter, and colloidal particles (e.g. filtering, gravity separation by coagulation, flocculation & clarification), removal or reduction of scale-forming minerals (e.g. lime softening, deionization and ion exchange), and removal of free fats, oil and grease (e.g. coagulation, flotation, and filtration). When metal ion leakage into the cathode is undesirably encountered, the following procedures or modifications may be considered: a purge or flush step of the cathode with deionized water, acid, base, chelating agent, or other cleaning solution, a potentiostatic cleaning procedure, a modification to the ion-exchange membrane to make it more selective for protons with respect to metallic cations, and/or a modification of the cathode catalyst layer and gas diffusion layer to make them more hydrophobic to facilitate cleaning. When sodium ion (Na⁺) ion leakage into the membrane is undesirably encountered, an in-situ ion exchange cleaning procedure may be performed. And, when sodium ion leakage into the cathode is undesirably encountered, as above a purge or flush step of the cathode with deionized water, acid, base or other cleaning solution may be used. In particular, a deionized water purge that results in formation of sodium hydroxide can provide a valuable by-product which can be recovered. And when oxygen leakage into the cathode is undesirably encountered, MnO₂ or other catalyst can be incorporated into the cathode gas diffusion layer and/or catalyst layer in order to decompose hydrogen peroxide. To provide for certain of the preceding cleaning processes, the cell and/or system may, at the cathode side, comprise a drain for cleaning solutions and a valve at the hydrogen gas outlet to prevent solution entering the gas line during cleaning. Drains may be incorporated generally which drain into the wastewater outlet or other general disposal. For clean in place capability, power would be turned off to the cell or cells, and a valve at the wastewater inlet employed to bypass the wastewater and to hook up a cleaning solution line. A valve at the exit may be employed in order to collect the cleaning solution. A similar process could be used on the hydrogen line.

One of ordinary skill in the art can be expected to appreciate the factors involved and to be able to determine what is adequate and how to adjust parameters such as flow rates, etc. accordingly. As shown in the Examples, model wastewater can be treated without fouling the cell electrodes. Oxygen evolution on the anode side due to water electrolysis as a side reaction can help keep the electrode free from any organic film buildup. However, in other situations, occasional cleanup of the electrodes may be required and accomplished by temporary cell reversals or other techniques known to those in the art.

The advantages of the present methods and systems are numerous. Primarily, they offer improved energy efficiency in the treatment of polluted wastewater. No solid waste or sludge is produced, nor toxic by-product gases which otherwise would need to be treated later. No catholyte is employed at the cathode, no fresh water is needed to generate hydrogen, and no waste is produced there. Thus, no additional chemicals need be added nor later removed to accomplish treatment. The system is versatile and can effectively treat effluents from industrial and municipal wastewaters and can mineralize many pollutants and microorganisms under the same operating conditions, thus combining organic pollutant removal and disinfection in a single step. Fundamentally, a wide operating range of temperatures, pressures, and variable effluent flow rates may be used. The system is scaleable and can be considered for treatment of wastewater quantities ranging from milliliters to millions of liters. The electrolytic cell components are suitable for low cost, high volume manufacturing processes and/or are already being mass produced. Along with low cost construction, operating costs and energy consumption are low, especially considering the possible capture of high purity by-product hydrogen for energy recovery, or use in other industrial operations.

Embodiments of the system may comprise multiple electrolytic cells in stacks, in either series and/or parallel flow arrangements, as shown in FIGS. 4 to 7. For example, wastewater can be split and supplied to multiple stacks of cells and the flows combined thereafter at the cell or stack outlets, as shown in FIGS. 6 and 7. Each of the stacks of electrolytic cells may be different, for example, operating at different operating conditions and/or comprising different components. This is particularly useful for wastewaters that comprise appreciable concentrations of oxidizable constituents such as ammonia, hydrogen sulfide, metals and inorganics, and/or constituents that decompose at low electrolytic potentials, decompose with heat, or decompose with different catalysts. By optimizing the design of each stack for particular types of contaminants that may be in the wastewater stream, improved contaminant removal, energy efficiency, operating costs and cell lifetime may be achieved. As shown in the Examples below, cell (and membrane electrode assembly) designs and operating conditions have a large effect on removal of different contaminants.

In FIG. 4, first cell stack 201 is connected upstream and in series with second cell stack 202, wherein the anode outlet from first cell stack 201 is connected to the anode inlet of second cell stack 202, with hydrogen generated being collected from all cathodes via stream 203. At least one of first cell stack 201 and second cell stack 202, operates at different operating conditions than the other cell stack, such as anolyte flow rate, current, voltage, temperature and/or pressure.

In FIG. 5, three stacks sharing end plates are connected in flow series such that reactant flows from first first stack 204, then to second stack 205, and then to third stack 206. Each of stacks 204, 205, 206 may be a monopolar design or a bipolar design, as described in the foregoing, and may comprise of one or more cells. Each of the stacks is sandwiched between shared end plates 207, 208, which may contain pockets (not shown) for each of the stacks for keeping them in place. The end plates may be electrically insulating or electrically conductive. If the end plates are electrically insulating, then current or voltage control may take place at each of the anode and cathode electrodes to provide different operating conditions to each of the cells. Alternatively, each stack may have conductive end plates (placed in the pocket of the insulating end plate), and current or voltage control may take place at the end plates so that operation of each cell in the stack will be the same, but may be different from other stacks. In yet another alternative, conductive end plates may be used and, current or voltage control may take place at the conductive end plates so that each of the stacks will operate at the same conditions. Furthermore, there may be a collection means between stacks for collecting wastewater for sampling and determining pollutant levels that can optionally be used for controlling current or voltage of the downstream segments (not shown). One skilled in the art will appreciate the advantages of such a design, such as requiring less hardware and balance of plant, while providing a smaller footprint. In addition, in a monopolar design, each of the cells can be operated at different current or voltage and, thus, the performance of each cell can be controlled to be the same without affecting other cells. (Electrical connections have been omitted from FIG. 5 for ease of understanding.) While that stacks are shown to be connected in flow series, it will be understood that the stacks may also be connected in different ways, such as that described for FIGS. 6 and 7.

In another embodiment, as shown in FIG. 6, first cell stack 210 is connected upstream and in series with second cell stack 211 and third cell stack 212, with hydrogen generated being collected from all cathodes via stream 213. In this embodiment, at least one of first cell stack 210, second cell stack 211, and third cell stack 212 operates at different operating conditions than the other cell stacks, such as anolyte flow rate, current, voltage, temperature and/or pressure. The anolyte flow from first cell stack 210 is divided among second and third cell stacks 211, 212 via valve 214. One skilled in the art will appreciate that the first stack may be in a reverse position than is shown in FIG. 5, that is, the second stack and the third stack are in a parallel configuration with the first stack in series downstream from the second and third stacks such that the two anolyte streams exiting from second and third stacks combine into a single anolyte stream that is then fed to the first stack.

In yet another embodiment, as shown in FIG. 7, first cell stack 210, second cell stack 211, and third cell stack 212 are connected in a parallel arrangement, such as the anolyte flows are divided among stacks 210, 211, and 212 via valve 215. In this embodiment, at least one of first cell stack 210, second cell stack 211, and third cell stack 212 operates at different operating conditions than the other cell stacks. Such an arrangement may be useful in situations where a portion of the wastewater is reused for other applications that can tolerate a higher contaminant level while another portion of the wastewater is recycled and requires a lower contaminant level. For example, a portion of the wastewater can be treated to reduce chemical oxygen demand (COD) or ammonia or hydrogen sulfide to about 99% removal level so that the treated wastewater can be reused in a reverse osmosis unit or other industrial process requiring a very low COD or other pollutants, while the remaining portion of wastewater only needs to be treated to a 50% COD level for discharge and/or for disinfection.

In further embodiments, treatment unit 218, which may be a filter, a degas unit, a pH controller, or the like, may be placed upstream of stacks, between stacks and/or downstream at the end of the stacks, as shown in FIGS. 5 and 6, or between stacks in FIG. 4 (treatment units 218 not shown). As a filter, treatment unit 218 captures oxidation and decomposition products between stacks to prevent such products from entering downstream stacks. As a filter, treatment unit 218 may be cleaned by any method known in the art, such as backwashing, air or gas scouring, and filter cartridge replacement. As a degas unit, treatment unit 218 removes product gases to prevent such products from entering downstream stacks. As a pH controller, treatment unit 218 adjusts wastewater pH to prevent corrosion of downstream components, precipitates pollutants and/or delivers effluent with requested pH.

For example, for wastewater with organic contaminants, ammonia, and hydrogen sulfide, two electrolytic cell stacks may be connected in series. The first cell stack may be set at a lower voltage and current and higher anolyte flow rate to oxidize the ammonia, hydrogen sulfide and other easily oxidized organics. The second cell stack may be set at a higher voltage and current, and at the same or lower anolyte flow rate to oxidize the remaining organics. A filter between the first and second cell stack may be employed to remove oxidation products, such as sulfur-containing oxidation product. Also, a degas unit may be placed in front of the first stack, between the first and second stack and/or at the end of the second stack. Also pH may be adjusted after the first stack and/or after the second stack.

In another example, for wastewater containing various molecular weight organic contaminants, such as volatile and petroleum hydrocarbons, a first cell stack may be operated at a lower voltage and/or current to oxidize the bulk of the easily oxidized organic constituents and partially oxidize the large molecular weight organics, and a second cell stack connected in series and operated at a higher voltage and/or current to oxidize the remaining species.

Additionally, or alternatively, each of the cell stacks may comprise different components, such as anode fluid delivery layers, anode filter layers, catalyst composition and loadings, flow field plate type and design, polymer electrolyte membranes, electrode active areas, and cell count. Again, each stack of electrolytic cells may be optimized for different types of contaminants that may be in the wastewater stream. For instance, as shown in the Examples (see Tables 1 through 9), certain anode catalysts have a higher affinity for oxidizing different contaminants. Therefore, in wastewater streams with several different contaminants, one skilled in the art will be able to determine the best anode catalyst (or membrane electrode assembly design) for a particular type of contaminant, and use them in separate groups of segmented electrolytic cells or stacks of electrolytic cells for improved contaminant removal from the wastewater stream.

The method can comprise a combination of stacks with filter layers and stacks without. The filter layers may be either electrically conductive or non-conductive. The anode and cathode flow field plates may also be electrically conductive or non conductive or a combination of both.

For influent wastewaters comprising solid particles, the method can comprise stacks with an anode filter layer adjacent the anode fluid diffusion layer and stacks without a filter layer. For instance, if particles are in the incoming wastewater, a first set of stacks may desirably have such filter layers. An intermediate filtration step may then be performed and the next stacks may not require such filter layers. For influent wastewaters without any particles, the first set of stacks may not have such filter layers but if solid particles are produced in the oxidation and/or decomposition process, the subsequent stacks amy desirably then have such filter layers. Also intermediate and post-filtration steps and/or pH adjustment steps and/or degas steps may also be employed.

The anode flow field plate may be electrically conductive with a conductive anode filter layer and fluid diffusion layer. Alternatively, for instance if the wastewater is very corrosive, the flow field plate may be made of plastic, the anode filter layer may be a non-conductive filter layer, while the fluid diffusion layer will be electrically conductive.

In one example, for landfill leachates, which typically contain ammonia nitrogen, organics, and chloride ions the wastewater treatment system may comprise three separate electrolytic cell stacks connected in series. The first cell stack is designed to oxidize ammonia at lower voltages, such as about 1.4 to about 2.0 V, while running the anolyte at a relatively faster flow rate than the second cell stack because of the faster reaction rate and/or lower contaminant concentration. The stack may also comprise fewer cells than the second cell stack because of the faster reaction rate. The reactions in the first electrolytic cell stack may include the following: 2NH₃+6OH⁻→N₂+6H₂O+6e ⁻  (1) 2Cl⁻→Cl₂+2e ⁻  (2) Cl₂+H₂O→HOCl+H⁺+Cl⁻  (3) ⅔NH₃+HOCl→⅓N₂+H₂O+H⁺+Cl⁻  (4)

The second electrolytic cell stack is designed to oxidize organics, and may operate at a higher voltage than the first electrolytic cell stack such as greater than about 2.5 V, while running the anolyte at a relatively slower flow rate than the other cell stack because of the slower reaction rates and/or higher concentration of contaminants Additionally, or alternatively, the stack may also comprise more cells than the other cell stacks because of the slower reaction rates and/or higher concentration of contaminants.

The third electrolytic cell stack is designed to remove the remaining microorganism contaminants from wastewater stream by disinfection. Therefore, the third cell stack may comprise means for generating free chlorine from chloride ions in chloride-containing wastewaters. The third cell stack may operate at a lower voltage than the second cell stack because the voltage for generating free chlorine is generally lower than the voltage required to generate hydroxyl radicals, and may comprise fewer cells than the second cell stack because less residence time is required for the wastewater since once free chlorine is generated, free chlorine will disinfect the wastewater and has residual disinfection.

One skilled in the art will recognize that the embodiments described above can be further combined in a number of different ways to optimize contaminant removal from the wastewater stream as well as improving energy efficiency and lifetime of the electrolytic cells. The embodiments described above merely illustrate several aspects of the invention and should not be construed to be limiting in any way.

In the construction of multiple cell systems, conductive layers may be employed between the fluid diffusion layers and plates or between the gas diffusion layers and plates. Alternatively, conductive foils or membranes may be welded to the fluid diffusion layers or gas diffusion layers.

The following examples are provided to illustrate certain aspects of the invention and particularly the various results obtained when employing electrolytic cells of the invention with different cell components and/or which are operated under different conditions. These examples however should not be construed as limiting in any way.

Examples

Numerous laboratory scale solid polymer electrolyte electrolytic cells were constructed as shown generally in FIG. 2 and were used to remove contaminants from wastewater samples via the method of the invention. The contaminants removed were either Acid Blue 29, phenol, acetaminophen, ibuprofen, Kraft mill effluent, or formic acid and these were present in different concentrations as indicated below.

The test electrolytic cells all employed a single membrane electrode assembly (MEA) comprising fluid and gas distribution layers adjacent to each of the anode and cathode electrodes. The fluid distribution layers were made of various porous carbon papers on which various microporous sublayers had been applied (as indicated below) and niobium mesh with a tungsten gauze sublayer. In some cases, commercially obtained MEAs were used and in other cases, catalyst layers comprising special catalyst compositions were prepared and applied to the fluid distribution layers (again as indicated below). The MEAs with fluid diffusion layers were clamped between graphite resin composite plates in which serpentine flow field channels had been machined. The size of the MEA varied somewhat from cell to cell as indicated below, but was of order of 50 cm² in size.

In these laboratory scale tests, several thicknesses of porous graphite paper from Toray were used as substrates for the fluid diffusion layers (i.e. Toray™ TGP-H-030=110 μm, TGP-H-60=190 μm, TGP-H-90=280 μm, TGP-H-120=370 μm). The papers were impregnated with PTFE using multiple successive conventional dip or flow techniques to build up the thickness of the PTFE coating slowly without forming cracks. Each coating layer was dried to remove water at 80° C. The PTFE impregnated substrate was either sintered at 400° C. for 10 minutes to increase the hydrophobicity of the surface before applying the microporous sublayer coatings, or was left unsintered to allow for controlled penetration of microporous coating solution.

Microporous sublayer coatings were then applied to the fluid diffusion layer substrates. Suspensions of electrically conductive particles and hydrophobic PTFE were prepared in solutions comprising water, wetting agent, and pore formers as indicated in Table 1 below. First, the electrically conducting particles were suspended in water and wetting agent by dispersing/mixing at 1500 rpm for 5 minutes. Then, the PTFE and pore former in water were added and mixed at 2500 rpm using a high shear mixer for 30 minutes or longer until no agglomeration is present (determined by fineness of grind gage). The sublayer suspension was then applied to the substrates either by rod or blade coating. The coated substrates were heated to remove water and then were calendared. Finally, both the wetting agent and pore former were removed and the applied PTFE was sintered by heating the coated substrates for 10 minutes at 400° C. Table 1 below summarizes the various sublayer compositions of the 8 different sublayers appearing in these Examples. Sublayer #s 4, 5, and 6 had the same composition and were made in the same manner but were applied in different amounts to the substrates involved.

TABLE 1 Sublayer Electrically conducting Hydrophobic Pore former and # particles PTFE rheology modifier Wetting Agent 1 5 wt. % Super P-Li ™ 2 wt. % 3 wt. % HPMC + 0.15 wt. % Tergitol ™ carbon black 90 wt. % H₂O 2 5 wt. % Timrex HSAG300 1 wt. % 1 wt. % HPMC + 0.2 wt. % Tergitol ™ graphite ™ 92.8 wt. % H₂O 3 2.5 wt. % Timrex KS150 ™ + 1 wt. % 1 wt. % HPMC + 0.6 wt. % Tergitol ™ 2.5 wt. % KS25 ™ graphite 92.4 wt. % H₂O 4, 5, 6 5 wt. % Timrex KS25 ™ 1 wt. % 1 wt. % HPMC + 0.4 wt. % Tergitol ™ graphite 92.6 wt. % H₂O 7 5.5 wt. % Timrex KS25 ™ 2 wt. % 1 wt. % HPMC + 0.5 wt. % Tergitol ™ graphite 91 wt. % H₂O 8 3.5 wt. % Timrex KS25 ™ 2 wt. % 1 wt. % HPMC + 0.5 wt. % Tergitol ™ graphite + 1.5 wt. % MnO₂ 91.5 wt. % H₂O 9 5 wt. % Niobium 1 wt. % 1 wt. % HPMC + 0.4 wt. % Tergitol ™ 92.6 wt. % H₂O Notes: Timrex HSAG300 ™ graphite has a particle size distribution with 90% <32 μm, and a surface area = 280 m²/g Super P-Li ™ conductive carbon black has 40 nm particle size and a surface area of 62 m²/g Timrex KS150 synthetic graphite has a particle size distribution with 95% <180 μm Timrex KS25 synthetic graphite has a particle size distribution with 90% <27.2 μm and a surface area of 12 m²/g MnO₂ powder has <5 μm particle size distribution HPMC stands for hydroxypropyl methylcellulose >95% of niobium was −325 mesh powder

Nine different anode catalyst layers (denoted A1 to A9) and five different cathode catalyst layers (denoted C1 to C5) appear in these Examples. The various catalyst layer and preparation suspension compositions are summarized in Table 2 below. A1 and C1 were commercially obtained platinum catalyst layers coated on a membrane electrolyte which were provided as a complete catalyst coated membrane (CCM) product from Ion Power, Inc. and thus do not appear in Table 2. The catalyst layers appearing in Table 2 were applied in the form of a suspension to the sublayer coated fluid diffusion layers or membrane electrolytes as indicated in Tables 4-7 below. The suspensions were prepared by adding the indicated catalyst and electrical conductor powder to a liquid carrier. The suspension was mixed at 2500-3500 rpm for about 30 minutes after which the proton conductor (electrolyte) was added and mixed further at 2500 rpm for 15 minutes. The catalyst coating suspension was then sparingly sprayed using multiple passes onto each surface of the membrane (CCM) or onto the fluid distribution layer and cathode gas diffusion layer (electrodes) using an air-powered, gravity-fed spray gun. The coating was dried between passes until the desired coating weight was reached.

TABLE 2 Catalyst Proton Layer Electrocatalyst Electron Conductor Liquid Carrier Conductor A2 3.0 wt. % ATO(1) 0.3 wt. % Silver + 0.3 40.0 wt. % none wt. % Super P-Li ™ Isopropanol + carbon black 56.4 wt. % H2O A3, A6 2.5 wt. % ATO(1) 0.25 wt. % Timrex 50 wt. % 1 wt. % HSAG300 ™ graphite Isopropanol + EW1100 53.75 wt. % Nafion ™ H2O A4 2.5 wt. % ATO(2) 0.25 wt. % Timrex 50 wt. % 1 wt. % HSAG300 ™ graphite Isopropanol + EW1100 53.75 wt. % Nafion ™ H2O A5 1.0 wt. % ATO(3) + 0.4 wt. % carbon support 20 wt. % 1 wt. % 0.25 wt. % Platinum and 1.5 wt. % HSAG300 ™ Isopropanol + EW1100 graphite support 75.85 wt. % Nafion ™ H2O A7 2.0 wt. % ATO(1) + 0.25 wt. % Sn—Ag 25 wt. % 1.0 wt. % 0.5 wt. % MnO2 + Isopropanol + EW1100 0.75 wt. % SnO2 70.5 wt. % H2O Nafion ™ A8 2.0 wt. % ATO(2) 0.5 wt. % Ta and 0.5 wt. % 25 wt. % 1.0 wt. % Nb and 0.5 wt. % TiC Isopropanol + EW1100 70.5 wt. % H2O Nafion ™ A9 1.0 wt. % ATO(2) + 0.25 wt. % Timrex 25 wt. % 1.0 wt. % 1.0 wt. % ATO(4) HSAG300 ™ graphite Isopropanol + EW1100 70.5 wt. % H2O Nafion ™ C2, C3, C4 1.5 wt. % Platinum 2.0 wt. % carbon support 25 wt. % 1.0 wt. % Isopropanol + EW1100 70.5 wt. % H2O Nafion ™ C5 1.5 wt. % Pt + 0.5 2.0 wt. % carbon support 20 wt. % 1.0 wt. % wt. % MnO2 Isopropanol + EW1100 75 wt. % H2O Nafion ™ Notes: ATO(1) stands for antimony tin oxide nanoparticles; ratio of Sb₂O₅:SnO₂ is 10:90 wt %; 22-44 nm particle size; and surface area of 20-40 m²/g ATO(2) was ATO(1) which had been heat treated for 4 hours at 550° C. in air ATO(3) was antimony tin oxide decorated Timrex HSAG300 ™ graphite ATO(4) was Nb and Sb doped tin oxide particles; Nb₂O₅:Sb₂O₅:SnO₂, nominal ratio 5:10:85 wt. % The platinum used was HiSPEC 4100 ™; nominally 40% by weight on carbon support Timrex HSAG300 ™ graphite is a conductive, high surface area graphite having a particle size distribution in which 90% <32 μm; and a surface area of 280 m²/g Super P-Li ™ was a conductive carbon black; with 40 nm particle size; and a surface area of 62 m²/g The silver used was a spherical powder, 99.9%(metals basis), having a particle size distribution of 1.3-3.2 μm; and a surface area of 0.3-0.7 m²/g MnO₂ powder had <5 μm particle size distribution Sn—Ag was an alloy nanopowder, with <150 nm particle size, 3.5% Ag SnO₂ was −325 mesh powder Nafion ™ EW1100 was a dispersion comprising colloidal particles in a 10 wt. % solution Ta was −325 mesh powder Nb was −325 mesh powder TiC powder had ≤4 μm particle size distribution

Further, in the above, the ATO(3) was prepared by dissolving 9.5 gm SnCl₂-2H₂O and 0.5 gm SbCl₃ in 10 ml concentrated HCl acid. The mixture was stirred until the solution was clear. 10 gm of pre-treated Timrex HSAG300™ graphite was then dispersed in 100 ml ethanol. This graphite suspension was heated to 80-90° C. and the acid solution was added slowly while continuing to stir. Heating and stirring continued until the ethanol evaporated. The powder product was filtered and washed with de-ionized water and then dried in an oven at 100° C. In this procedure the Timrex HSAG300™ had been pre-treated by first combining 0.25 gm PdCl₂, 12.5 gm SnCl₂-2H₂O, 150 ml de-ionized water, and 75 ml concentrated HCl acid, stirring at room temperature until green in colour (≥1 hr), then adding 20 gm of the graphite powder to this suspension, stirring for 1-3 minutes, and finally filtering, rinsing and drying the powder.

The test MEAs comprising the fluid diffusion layers were bonded together into unitary assemblies before testing. When employing commercially obtained and in-house manufactured catalyst coated membrane electrolytes, these were placed between an appropriate anode fluid distribution layer and cathode fluid distribution layer (henceforth referred to as cathode gas diffusion layer because the fluid at the cathode side was always gaseous) and were either hot pressed at 140° C. for 5 minutes or left un-bonded for testing. When employing the catalyst coated fluid diffusion layers described herein, these electrodes were placed on either side of a commercially obtained membrane electrolyte and hot pressed at 140° C. for 5 minutes to bond them together. PTFE tape was used to mask the edges of un-bonded CCMs to provide a dimensionally stable perimeter for the cell assembly.

The compositions and loadings of the various catalyst layers and fluid distribution layers used in the MEAs in these Examples are summarized in Table 3 below.

TABLE 3 MEA Anode fluid distribution layer Anode Catalyst Layer³ Membrane Cathode catalyst layer^(b) Cathode gas diffusion layer A Substrate: TGP60 + 10 wt. % #A1 - 0.3 mg/cm² Pt Nafion ™ #C1 - 0.3 mg/cm² Pt Substrate: TGP60 + 10 wt. % PTFE XL100 PTFE Microporous layer #1: 6 gm/m² Microporous layer #1: 6 gm/m² carbon black + 25 wt. % PTFE carbon black + 25 wt. % PTFE B Substrate: TGP90 + 10 wt. % #A1 - 0.3 mg/cm² Pt Nafion ™ #C1 - 0.3 mg/cm² Pt Substrate: TGP90 + 10 wt. % PTFE XL100 PTFE Microporous layer #2: 6.5 gm/m² Microporous layer #2: 6.5 gm/m² graphite − 15 wt. % PTFE graphite + 15 wt. % PTFE C Substrate: TGP90 + 10 wt. % #A2 - 4.5 mg/cm² Nafion ™ #C2 - 1 mg/cm² Pt − Substrate: TGP120 + 10 wt. % PTFE ATO(1) − 25 wt. % Ag − N211 30 wt. % Nafion ™ PTFE Microporous layer #3: 10 gm/m² 20 wt. % PTFE Microporous layer #3: 10 gm/m² graphite − 15 wt. % PTFE graphite + 15 wt. % PTFE D Substrate: TGP60 + 10 wt. % #A3 - 2.5 mg/cm² Nafion ™ #C2 - 1 mg/cm² Pt − Substrate: TGP60 + 10 wt. % PTFE ATO(1) − 10 wt. % N211 30 wt. % Nafion ™ PTFE Microporous layer #4: 20 gm/m² HSAG300 − 30 wt. % Microporous layer #4: 20 gm/m² graphite − 15 wt. % PTFE Nafion ™ graphite + 15 wt. % PTFE E Substrate: TGP60 + 10 wt. % #A4 - 2.5 mg/cm² Nafion ™ #C3 - 1.5 mg/cm² Pt + Substrate: TGP60 + 10 wt. % PTFE ATO(2) − 10 wt. % XL100 30 wt. % Nafion ™ PTFE Microporous layer #4: 20 gm/m² HSAG300 + Microporous layer #4: 20 gm/m² graphite − 15 wt. % PTFE 30 wt. % Nafion ™ graphite + 15 wt. % PTFE F Substrate: TGP120 + 10 wt. % #A4 - 2.5 mg/cm² Nafion ™ #C1 - 0.3 mg/cm² Pt Substrate: TGP120 + 20 wt. % PTFE ATO(2) − 10 wt. % XL100 PTFE Microporous layer #3: 10 HSAG300 + 30 wt. % Microporous layer #6: 20 gm/m² gm/m² graphite − 15 wt. % PTFE Nafion ™ graphite + 30 wt. % PTFE F1 Substrate: TGP120 + 10 wt. % #A4 - 2.5 mg/cm² Nafion ™ #C1 - 0.3 mg/cm² Pt Substrate: TGP120 + 20 wt. % PTFE ATO(2) − 10 wt. % XL100 PTFE Microporous layer #3: 10 gm/m² HSAG300 + 30 wt. % Microporous layer #8: 15 gm/m² graphite − 15 wt. % PTFE Nafion ™ graphite + 5 g/m² MnO₂ + 30 wt. % PTFE G Microporous layer #3: 10 gm/m² #A5 - 2.5 mg/cm² Nafion ™ #C4 - 1.5 mg/cm² Pt + Microporous layer #6: 20 gm/m² graphite − 15 wt. % PTFE ATO(1) − 20 wt. % Pt − XL100 30 wt. % Nafion ™ graphite + 30 wt. % PTFE Microporous layer #3: 10 gm/m² 30 wt. % Nafion ™ Microporous layer #3: 10 gm/m² graphite − 15 wt. % PTFE graphite + 15 wt. % PTFE H Substrate: TGP60 + 10 wt. % #A4 - 2.5 mg/cm² Nafion ™ #C4 - 1.5 mg/cm² Pt + Substrate: TGP60 + 10 wt. % PTFE ATO(2) − 10 wt. % NR211 30 wt. % Nafion ™ PTFE Microporous layer #3: 10 gm/m² HSAG300 + 30 wt. % Microporous layer #6: 20 gm/m² graphite − 15 wt. % PTFE Nafion ™ graphite + 15 wt. % PTFE I Substrate: TGP30 + 10 wt. % #A3 - 2.5 mg/cm² Nafion ™ #C1 - 0.3 mg/cm² Pt Substrate: TGP30 + 10 wt. % PTFE ATO(1) − 10 wt. % XL100 PTFE Microporous layer #3: 10 gm/m² HSAG300 − 30 wt. % Microporous layer #3: 10 gm/m² graphite − 15 wt. % PTFE Nafion ™ graphite + 15 wt. % PTFE J Substrate: TGP60 + 10 wt. % #A5 - 2.5 mg/cm² Nafion ™ #C4 - 1.5 mg/cm² Pt + Substrate: TGP60 + 40 wt. % PTFE ATO(3) − 20 wt. % Pt − N211 30 wt. % Nafion ™ PTFE Microporous layer #5: 15 gm/m² 30 wt. % Nafion ™ Microporous layer #7: 10 gm/m² graphite − 15 wt. % PTFE graphite + 30 wt. % PTFE K. Substrate: TGP120 + 10 wt. % #A6 - 5 mg/cm² Nafion ™ #C4 - 1.5 mg/cm² Pt + Substrate: TGP120 + 10 wt. % K2 PTFE ATO(1) − 10 wt. % XL100 30 wt. % Nafion ™ PTFE Microporous layer #3: 10 gm/m² HSAG300 − 30 wt. % Microporous layer #4: 20 gm/m² graphite − 15 wt. % PTFE Nafion ™ graphite + 15 wt. % PTFE L Substrate: TGP120 + 10 wt. % #A7 - 1.1 mg/cm² Nafion ™ #C2 - 1 mg/cm² Pt − Substrate: TGP120 + 30 wt. % PTFE ATO(1) − 0.25 mg/cm² XL100 30 wt. % Nafion ™ PTFE Microporous layer #3: 10 gm/m² MnO₂ + 25 wt. % SnO₂ − Microporous layer #7: 10 gm/m² graphite − 15 wt. % PTFE 10 wt. % Sn—Ag + graphite + 30 wt. % PTFE 30 wt. % Nafion ™ M Nb gauze − 40 mesh 17.8 cm dia. #A7 - 1.1 mg/cm² Nafion ™ #C2 - 1 mg/cm² Pt − Substrate: TGP120 + 30 wt. % wire − ATO(1) − 0.25 mg/cm² XL100 30 wt. % Nafion ™ PTFE W gauze − 100 mesh 2.54 cm dia. MnO₂ + 25 wt. % SnO₂ − Microporous layer #7: 10 gm/m² wire 10 wt. % Sn—Ag + graphite + 30 wt. % PTFE 30 wt. % Nafion ™ N Substrate: TGP120 + 10 wt. % #A8 - 0.6 mg/cm² Nafion ™ #C2 - 1 mg/cm² Pt − Substrate: TGP120 + 30 wt. % PTFE ATO(2) − 0.15 mg/cm² XL100 30 wt. % Nafion ™ PTFE Microporous layer #3: 10 gm/m² Nb + 0.15 mg/cm² Ta + Microporous layer #8: 15 gm/m² graphite − 15 wt. % PTFE 0.15 mg/cm² TiC graphite + 5 g/m² MnO₂ + 30 wt. % PTFE O Substrate: Graphite felt #A7 - 1.1 mg/cm² Nafion ™ #C2 - 1 mg/cm² Pt − Substrate: TGP60 + 30 wt. % ATO(1) − 0.25 mg/cm² XL100 30 wt. % Nafion ™ PTFE No Microporous layer MnO₂ + 25 wt. % SnO₂ − Microporous layer #4: 20 gm/m² 10 wt. % Sn—Ag + graphite + 15 wt. % PTFE 30 wt. % Nafion ™ P Substrate: Niobium screen #A9 - 2.5 mg/cm³ Nafion ™ #C2 - 1 mg/cm² Pt − Substrate: TGP60 + 30 wt. % ATO(2) − 2.5 mg/cm² XL100 30 wt. % Nafion ™ PTFE Microporous layer #9: 10 gm/m² ATO(4) − 10 wt. % Microporous layer #4: 20 gm/m² niobium + 15 wt. % PTFE HSAG300 − 30 wt. % graphite + 15 wt. % PTFE Nafion ™ Q Substrate: TGP120 + 20 wt. % #A9 - 2.5 mg/cm² Nafion ™ #C2 - 1 mg/cm² Pt − Substrate: TGP120 + 20 wt. % PTFE ATO(2) − 2.5 mg/cm² XL100 30 wt. % Nafion ™ PTFE Microporous layer #7: 10 gm/m² ATO(4) − 10 wt. % Microporous layer #7: 10 gm/m² graphite − 30 wt. % PTFE HSAG300 − 30 wt. % graphite + 30 wt. % PTFE Nafion ™ ¹Unless otherwise indicated, the substrate with PTFE was not sintered before sublayer coating

The electrochemical cell assembly was completed by sandwiching the test MEAs between anode and cathode flow field plates made of polymer-graphite composite. A 4 pass serpentine channel had been machined in the cathode flow field plate with a 1 mm channel width, 1 mm channel height, 1 mm landing width and a geometric area of 50 cm². Two different anode flow field plates were used; the first having a 4 pass serpentine channel machined in the flow field plate with a 1 mm channel width, 1 mm channel height, 1 mm landing width and a geometric area of 50 cm², and the second having a single channel machined therein with a 5 mm channel width, 8 mm channel height, 2 mm landing width, and a geometric area of 50 cm². A spiral in-line mixing component, manufactured from twisted PTFE tape, 2 mm in width, was used with the single channel anode flow field plate and the channel interior was coated with PTFE. The sealing gaskets used were made of Viton® and Gore®, the current collectors were gold coated copper, and the end compression plates were made of steel and contained interior electrical resistance heating elements. In all the experimental tests below, the 4 pass channel design was used except for the test involving MEA K2 in Table 5 which used the single channel and in-line mixing component.

Testing then involved preparing model contaminated wastewaters (>1 L of solution) with the specified pollutant in de-ionized water. The electrochemical cell temperature was kept constant using the internal resistive heating elements, a temperature controller, and thermocouple. Several test temperatures were used as indicated below. Wastewater comprising the indicated contaminant was then flowed through the anode of the test cell using a peristaltic pump at a rate of 270 mL/hour while a constant DC voltage was applied to the current collectors. The valve downstream from the anode exhaust was used in selected trials to provide pressurized flow. The cathode inlet of the test cell was sealed and the cathode exhaust was also provided with a valve downstream to provide slightly pressurized hydrogen gas exhaust. The majority of tests were run at atmospheric pressure at the anode exhaust and slight pressure (<1 psi) at the cathode exhaust as a result of filling the hydrogen storage container. No water or purge gases were used or required on the cathode. No supporting electrolyte of any kind was used at the cathode in any test. The wastewater effluent was collected in a plastic jug and the product gases were released to the atmosphere.

Tables 4, 5, 6, 7, 8 and 9 below summarize the results obtained for the tests involving Acid Blue 29 dye, phenol, acetaminophen, formic acid, ibuprofen, and Kraft effluent respectively.

In the case of the Acid Blue 29 dye pollutant, colour measurements were used to quantify the efficacy of treatment. The % of colour removal was determined with a UV/VIS Spectrophotometer by comparing absorbance against samples of known concentrations.

In the case of the other pollutants tested, the chemical oxygen demand (COD) was used to quantify the efficacy of treatment. COD is used as a measurement of pollutants in wastewaters and natural waters. Both organic and inorganic components of a sample are subject to oxidation, but in most cases the organic component predominates and is of the greatest interest (ref. Standard Methods for the Examination of Water and Wastewater, 21^(st) Edition, APHA, AWWA, WEF, ©2005). In general, the oxidation of specific compounds is characterized by the extent of degradation of the final oxidation products (ref: Industrial Water Quality, 4th edition, W. Wesley Eckenfelder, Jr., Davis L. Ford and Andrew J. Englande, Jr. McGraw-Hill Companies, Inc. © 2009). The reason for this is that the degradation of the pollutant can be referred to in several ways. There is: (1) Primary degradation which involves a structural change in the parent compound; (2) Acceptable degradation (defusing) which involves a structural change in the parent compound to the extent that toxicity is reduced; (3) Ultimate degradation (mineralization) which involves conversion of organic carbon to inorganic CO₂; and (4) Unacceptable degradation (fusing) which involves a structural change in the parent compound resulting in an increase in toxicity. Any degradation process that does not lead to total mineralization of the organic constituents may potentially form end products that can be more toxic than the original compounds. FIG. 4 is a prior art illustration of how the change in original compound concentration can differ from that of the COD over the course of oxidation for refractory organic compounds such as phenol. Although at point A, the amount of original/parent compound has decreased to zero, the COD of the wastewater does not meet discharge limit for COD concentration.

Therefore, to quantify the pollutant removal efficacy of the system/process, ultimate degradation (mineralization) of the organic compounds is preferably measured by the chemical oxygen demand (COD). COD will report virtually all organic compounds, and is used for monitoring and control of discharges in industrial applications, discharge permits, and for assessing treatment plant performance COD is a measure of the total quantity of oxidizable components in a sample (e.g. carbon, hydrogen from hydrocarbons, nitrogen, sulfur, and phosphorus) and was measured here by Method 5220 C (EPA approved—Standard Methods for the Examination of Water and Wastewater, 21^(st) edition).

Samples of the treated wastewater were taken throughout the test periods and average values for colour and COD were determined in accordance with the pollutant present. The current across the test cells was generally stable and the average current density was also determined as reported below.

Tables 4 to 9 also list the energy consumption (the product of voltage, average current, and time over all the passes through the cell) per unit volume of wastewater. Where appropriate, the specific energy consumption per unit mass of COD mineralized is also listed.

Further, the hydrogen gas volume produced was measured in each case at the storage device. And from this, the efficiency of H₂ electrolysis was determined and listed in the Tables. Under ideal circumstances it requires 39.4 kWh of electricity at normal conditions (25° C. and 1 atm) to make 1 kg of hydrogen. This represents the higher heating value (HHV) of hydrogen, which includes the total amount of energy (thermal and electrical) to disassociate water at normal conditions. System efficiency is calculated by dividing the heating value (HHV) by the real energy input in units of kWh/kg. Industrial electrolyzer efficiencies generally are in the range of 52% to 82% (HHV).

TABLE 4 Colour Removal Efficiency Membrane Current Pollutant Energy Hydrogen of H₂ electrode Temp. Voltage density Removal % Consumption Generation electrolysis Wastewater Composition assembly (MEA)¹ (° C.) (V) (mA/cm²) colour (kWh/m³ ww) Rate (ml/hr) (HHV) 60 mg/l Acid blue 29 dye A 25 1.8 6.5 95 8 85 55.4 50 mg/l Acid blue 29 dye C 50 2.1 4 95 5 75 62.9 50 mg/l Acid blue 29 dye C 50 2.1 4 100 11 75 62.9 50 mg/l Acid blue 29 dye D 50 2.6 8 100 15 225 76.2 50 mg/l Acid blue 29 dye G* 50 2.3 6 95 5 115 58.7 50 mg/l Acid blue 29 dye F* 40 2.3 3 100 12 45 57.4 50 mg/l Acid blue 29 dye I 35 2.3 2 95 7 25 47.8 50 mg/l Acid blue 29 dye I 35 2.3 2 100 11 25 47.8 100 mg/l Acid blue 29 dye  P 40 5 0.5 90 2 5 35.2 Note: In Tables 4 to 9. * indicate that the catalyst layer was coated onto fluid and gas distribution layers: all the other MEAs comprise catalyst layers coated onto the membrane. All CCM based MEAs were tested unbounded while the others were tested bonded.

TABLE 5 Phenol Removal Chemical Membrane Energy Specific Efficiency Nominal Oxygen electrode Current Pollutant Consumption Energy Hydrogen of H₂ Wastewater Demand assembly Temp. Voltage density Removal (kWh/m³ Consumption Generation electrolysis Composition (COD mg/L) (MEA) (° C.) (V) (mA/cm²) % COD ww) (kWh/kg COD) Rate (ml/hr) (HHV) 500 mg/l phenol 1227 B* 50 1.8 6.5 39 16 32.4 not — measured 330 mg/l phenol 955 E* 35 2.3 2.5 72 19 27.6 30 57.4 500 mg/l phenol 1258 F 35 2 3 65 12 14.8 45 66.0 500 mg/l phenol 1258 F 40 2.1 4 85 17 16.3 55 61.5 500 mg/l phenol 1258 F1 40 2.2 4 80 19 18.9 60 64.0 330 mg/l phenol 955 F* 35 2.3 4.5 80 16 21.1 65 56.8 1100 mg/l phenol  2645 F* 40 2.2 4 80 38 18.6 55 58.7 2000 mg/l phenol  5266 F 30 2.3 2.5 72 42 11.2 30 45.9 500 mg/l phenol 1209 G* 25 2.7 4.5 40 18 37.1 75 48.9 250 mg/l phenol 578 J 35 2.8 4 70 15.5 38.3 60 50.3 250 mg/l phenol 578 K2 25 2.8 5 77 21 46.4 75 47.1 1000 mg/l phenol  2326 K 25 2.8 5 65 41 27.8 80 50.3 500 mg/l phenol 1130 L 35 2.3 1.5 70 9 10.9 15 45.9 500 mg/l phenol 1148 N 40 2.8 2 85 8.3 8.8 35 55.0 500 mg/l phenol 1149 O 35 2.8 2 93 8.3 10 35 55.0 250 mg/l phenol 596 O 30 2.8 1.5 95 4.8 8.5 20 50.3

TABLE 6 Acetaminophen Removal Chemical Membrane Efficiency Nominal Oxygen electrode Current Pollutant Energy Specific Energy Hydrogen of H₂ Wastewater Demand assembly Temp. Voltage density Removal Consumption Consumption Generation electrolysis Composition (COD mg/L) (MEA) (° C.) (V) (mA/cm²) % COD (kWh/m³ ww) (kWh/kg COD) Rate (ml/hr) (HHV) 500 mg/l 1000 E 35 2.7 3 89 21 23.6 35 57.0 acetaminophen 1 g/l 1778 H* 35 2.25 4 80 75 52.5 70 54.8 acetaminophen

TABLE 7 Formic Acid Removal Chemical Membrane Specific Nominal Oxygen electrode Current Pollutant Energy Energy Hydrogen Efficiency of H2 Wastewater Demand assembly Temp. Voltage density Removal Consumption Consumption Generation electrolysis Composition (COD mg/L) (MEA) (° C.) (V) (mA/cm²) % COD (kWh/m³ ww) (kWh/kg COD) Rate (ml/hr) (HHV) 2 ml/L 841 M 35 2.8 2.5 85 10 18 35 44.0 Formic acid 2 ml/L 841 M 35 2.8 2.5 95 15 24 35 44.0 Formic acid

TABLE 8 Ibuprofen Removal Chemical Membrane Specific Nominal Oxygen electrode Current Pollutant Energy Energy Hydrogen Efficiency of H₂ Wastewater Demand assembly Temp. Voltage density Removal Consumption Consumption Generation electrolysis Composition (COD mg/L) (MEA) (° C.) (V) (mA/cm²) % COD (kWh/m³ ww) (kWh/kg COD) Rate (ml/hr) (HHV) 0.1 g/l 383 Q 40 2.8 2 80 4.3 14.4 35 55.0 ibuprofen

TABLE 9 Kraft effluent removal Chemical Membrane Efficiency Nominal Oxygen electrode Current Pollutant Energy Specific Energy Hydrogen of H2 Wastewater Demand assembly Temp. Voltage density Removal Consumption Consumption Generation electrolysis Composition (COD mg/L) (MEA) (° C.) (V) (mA/cm²) % COD (kWh/m³ ww) (kWh/kg COD) Rate (ml/hr) (HHV) Kraft pulp 471 Q 40 2.8 2 60 5.2 19.2 35 44.0 & paper mill effluent after biological reactor

The results using these laboratory test cells show that electrochemical cells with non-liquid, polymer electrolytes, that contain no other added chemicals, and comprising low cost catalysts and other electrode components can provide equal or better removal efficiency as comparative prior art systems for recalcitrant Acid Blue 29 dye, phenol, acetaminophen, formic acid, ibuprofen, and Kraft pulp and paper mill effluent. In particular, these results can be obtained with substantially lower energy inputs (i.e. at current densities less than about 10 mA/cm² and applied voltages less than about 3 V), in some instances with greater than 60% energy reduction at 80% COD removal, with greater than 80% energy reduction at 95% COD removal and this is without including recoverable energy contributions from the hydrogen produced. A 20% increase in current efficiency was observed for Acid Blue dye 29, and over 60% increase for phenol and acetaminophen. Certain specific in-house prepared catalyst choices and electrode designs can lead to >40% improvement in performance.

Further still however, the inventive method efficiently produces hydrogen at a purity equivalent to commercial electrolyzers and in sufficient amounts such that an estimated additional 15-35% reduction in net energy consumption may be achieved depending on wastewater composition (assuming conversion of hydrogen back to electricity using a fuel cell stack operating at 50% efficiency and assuming 95% of the hydrogen was recovered). For illustrative purposes, FIG. 5 shows the average actual hydrogen generated from a number of tests performed at several different currents on phenol contaminated wastewater compared to ideal or perfect hydrogen generation. As can be seen, there is a high conversion of phenol contaminant to hydrogen.

In addition, the recoverable energy in a realistic scaled industrial system can be estimated based on the above. Assuming state-of the art fuel cells are used to convert the generated hydrogen back into electricity at 50% efficiency, Table 10 shows the expected recoverable energy in an industrial system operating as per the three data points shown in FIG. 5 above. In this Table, the system has been scaled up to treat 1 m³/hr 500 mg/l phenol wastewater, and it is assumed that the hydrogen generated is converted back to electricity with 95% utilization using 5 kW fuel cells operating at 50% efficiency.

TABLE 10 Hydrogen generation rate Recoverable Energy Operating conditions (m³/hr H₂) (kWh/m³ wastewater) 1^(st) data point in FIG. 5 6.7 12.5 2^(nd) data point 6.9 12.9 3^(rd) data point 8.5 15.9

All of the above mentioned U.S. patents and applications, foreign patents and applications and non-patent publications referred to in this specification, are incorporated herein by reference in their entirety.

While particular embodiments, aspects, and applications of the present invention have been shown and described, it is understood by those skilled in the art, that the invention is not limited thereto. Many modifications or alterations may be made by those skilled in the art without departing from the spirit and scope of the present disclosure. The invention should therefore be construed in accordance with the following claims. 

What is claimed is:
 1. A method for the energy efficient treatment of polluted wastewater comprising: providing at least a first and second solid polymer electrolyte electrolytic cell stack, the cells in the first and second stacks each comprising an anode comprising an anode catalyst layer and the anode catalyst layer comprising an anode catalyst, a cathode comprising a cathode catalyst layer and the cathode catalyst layer comprising a cathode catalyst wherein the cathode is liquid-electrolyte free, and a solid polymer membrane electrolyte separating the anode and the cathode; supplying a flow of wastewater comprising a pollutant to the anodes of the cells of the first and second electrolytic cell stacks without an added supporting electrolyte at a flow rate and flow pressure; providing a voltage less than about 3 volts across each of the cells in the first and second electrolytic cell stacks wherein the anode is positive with respect to the cathode; operating each of the cells in the electrolytic cell stacks at an operating temperature and a current density less than about 20 mA/cm2, thereby degrading the pollutant and generating hydrogen gas at the cathode; and exhausting the generated hydrogen gas from the cathode; wherein at least one stack component in the first solid polymer electrolyte electrolytic cell stack is different from the corresponding stack component in the second solid polymer electrolyte electrolytic cell stack; wherein the different stack component is selected from the group consisting of an anode fluid delivery layer, the anode catalyst, the anode catalyst layer, an anode flow field plate, an anode filter layer, the solid polymer electrolyte membrane, and the number of cells in the stack; and wherein the different stack component is the anode catalyst layer, and a catalyst active area and/or a catalyst loading in the anode catalyst layer of at least one of the cells of the first solid polymer electrolyte electrolytic cell stack differs by more than about 5% from a corresponding catalyst loading and/or catalyst active area respectively in the anode catalyst layer of at least one of the cells of the second solid polymer electrolyte electrolytic cell stack.
 2. The method of claim 1 wherein at least one of the catalyst loading and/or the catalyst active area in the anode catalyst layer of at least one of the cells of the first solid polymer electrolyte electrolytic cell stack differs by more than about 10% from the corresponding catalyst loading and/or the catalyst active area respectively in the anode catalyst layer of at least one of the cells of the second solid polymer electrolyte electrolytic cell stack.
 3. The method of claim 1 wherein an operating condition of the first solid polymer electrolyte electrolytic cell stack is different from the corresponding operating condition of the second solid polymer electrolyte electrolytic cell stack.
 4. The method of claim 3 wherein the different operating condition is selected from the group consisting of the flow rate of the wastewater, the flow pressure of the wastewater, the voltage, the operating temperature (Celsius), and the current density.
 5. The method of claim 4 wherein the different operating condition of the first solid polymer electrolyte electrolytic cell stack differs by more than about 5% from that of the second solid polymer electrolyte electrolytic cell stack.
 6. The method of claim 5 wherein the different operating condition of the first solid polymer electrolyte electrolytic cell stack differs by more than about 10% from that of the second solid polymer electrolyte electrolytic cell stack.
 7. The method of claim 1 comprising operating each of the cells in the electrolytic cell stacks at a current density less than about 10 mA/cm².
 8. The method of claim 1 wherein the cathode of each of the cells in the electrolytic cell stacks comprises no liquid catholyte nor liquid supporting electrolyte.
 9. The method of claim 1 wherein each of the first and second solid polymer electrolyte electrolytic cell stacks comprises a single electrolytic cell.
 10. The method of claim 1 wherein each of the first and second solid polymer electrolyte electrolytic cell stacks comprises more than one electrolytic cell.
 11. The method of claim 1 wherein the first solid polymer electrolyte electrolytic cell stack is connected upstream and in series flow with the second solid polymer electrolyte electrolytic cell stack and an anode outlet from the first solid polymer electrolyte electrolytic cell stack is connected to an anode inlet of the second solid polymer electrolyte electrolytic cell stack.
 12. The method of claim 11 wherein the first and second solid polymer electrolyte electrolytic cell stacks comprise common end plates.
 13. The method of claim 1 wherein the first solid polymer electrolyte electrolytic cell stack is connected in parallel flow with the second solid polymer electrolyte electrolytic cell stack and wherein the supplied wastewater is divided between the anode inlets of the first and second solid polymer electrolyte electrolytic cell stacks.
 14. The method of claim 1 comprising incorporating a treatment unit in the flow of wastewater.
 15. The method of claim 14 wherein the treatment unit is selected from the group consisting of a filter, a degas unit, and a pH controller.
 16. The method of claim 14 comprising incorporating the treatment unit upstream or downstream of either the first or the second solid polymer electrolyte electrolytic cell stack. 